Improvements on the design of carbon dioxide conversion to ... · Improvements on the design of carbon dioxide conversion to methanol process using Aspen ... the design of carbon

Improvements on the design of carbon dioxide conversion

to methanol process using Aspen Plus® interface

Arthur Vanhove

Dissertation to obtain the Master of Science Degree in

Chemical Engineering

Supervisor: Prof. Henrique Aníbal Santos de Matos

Examination Committee

Chairperson: Prof. Sebastião Manuel Tavares Silva Alves

Supervisor: Prof. Henrique Aníbal Santos de Matos

Members of the Committee: Prof. Maria Rosinda Costa Ismael

June 2015

I

Acknowledgments

Acknowledgments

I am using this opportunity to express my gratitude to everyone who supported me throughout

this thesis research project. I am thankful for their aspiring guidance, invaluably constructive criticism

and friendly advice during the project work. I am sincerely grateful to them for sharing their truthful and

illuminating views on a number of issues related to the project.

First and foremost, I thank my academic advisor, Mr. Prof. Dr. Henrique Matos, for accepting

me into his group. I express my warm thanks to Mr. Prof. Dr. Henrique Matos for his support and

guidance at Instituto Superior Técnico.

I would also like to thank my project colleagues, Pedro Pereira and Catarina Braz for the

enormous help and guidance throughout this project. Gratitude to Instituto Superior Técnico and all

the people who provided me with the facilities being required and conductive conditions for my thesis.

I would also like to extend my deepest gratitude to my family. Without their encouragement, I

would not have a chance to be in Lisbon. Also my best friends for visiting and supporting me during

this period.

Thank you,

Arthur Vanhove

III

Abstract

There has been a large increase in anthropogenic emissions of carbon dioxide (CO2) over the

past century due the industry. Global warming is amongst the effects of this greenhouse gas. Cimpor,

a Portuguese Cement plant, is well aware of this issue, as the cement industry is currently one of the

world’s main industrial sources of CO2 emissions. Many industries are implementing large-scale CO2

capture, usage and storage (CCUS). Although there are a lot of possibilities where CCUS can be

employed, estimated costs for its application appear to be extremely high; from today’s view, cement

production costs would roughly double.

The use of captured CO2 can become a profitable business, in addition to controlling CO2

concentration in the atmosphere. Among all the possible CO2-usage technologies, a CO2 conversion

to methanol process was simulated and improved with Aspen Plus®. The simulation for producing fuel

grade methanol is divided in sections, namely a gas compression section, a reaction section where

hydrogenation takes place, followed by a separation downstream section. Several improvements were

done based on utility optimization and heat integration between all process sections. Activated

analysis tools were used to compute the process parameters. Heat exchangers were sized and

summarized in TEMA sheets.

To argue the feasibility of CO2 conversion processes, nine different improvement scenarios on

the design of carbon dioxide conversion to methanol process were made. These improvements

resulted in a decrease of some Key Performance Indicators (KPI) between 50-70%, leading to a 69%

KPI enhancement of hydrogen efficiency and generating methanol at 1.6 thousand euro per tonne.

Keywords

Process modelling; Simulation CO2 usage; CO2 to methanol; Aspen Plus®; KPI

V

Resumo

Ao longo do último século tem-se verificado um aumento significativo das emissões de dióxido

de carbono (CO2) de origem antropogénica, resultante de uma intensificação da atividade industrial. O

aquecimento global está dentro dos efeitos da libertação deste gás de estufa. A Cimpor, uma

produtora de cimentos portuguesa, tem grande atenção a este problema, uma vez que a indústria

cimenteira é atualmente a maior fonte industrial de emissão de CO2 no mundo. Muitas indústrias

estão a implementar processos de captura, uso e armazenamento de CO2 (CO2 capture, usage and

storage – CCUS) em grande escala. Embora se possam usar técnicas de CCUS em várias

aplicações, os custos estimados aparentam ser demasiado elevados; no panorama atual, os custos

na indústria cimenteira tenderiam para o dobro.

O uso do CO2 capturado pode se tornar lucrativo, aliado ao controlo da libertação de CO2 para

a atmosfera. Dentro de todas as tecnologias possíveis para o uso de CO2, escolheu-se o processo de

conversão deste para metanol para ser simulado e otimizado em Aspen Plus®. A simulação para a

produção de metanol para ser usado como combustível é dividida por etapas, nomeadamente, uma

secção de compressão de gás, uma zona reacional de hidrogenação, seguida de uma secção de

separação a jusante do processo. Foram feitas várias melhorias ao processo, baseadas na

otimização do uso das utilidades e por integração energética entre todas as zonas do processo.

Foram também usadas ferramentas de análise para determinar os parâmetros do processo. Os

permutadores de calor foram dimensionados e as suas características apresentadas em folhas

TEMA.

Para verificar a viabilidade dos processos de conversão de CO2, foram testados nove cenários

diferentes de melhorias ao processo de conversão de CO2 a metanol. Estas melhorias resultaram na

diminuição dos Indicadores de Performance-Chave (Key Performance Indicators – KPI) entre 50-70%,

levando a um aumento de 69% de KPI na eficiência do hidrogénio e à produção de metanol a 1.6 mil

euros por tonelada.

Palavras-chave

Modelação de processos; Simulação do uso de CO2; CO2 para metanol; Aspen Plus®; KPI

VII

Contents

Acknowledgments ..................................................................................................................... I

Abstract ...................................................................................................................................III

Resumo ................................................................................................................................... V

Contents ................................................................................................................................ VII

List of tables ............................................................................................................................ IX

List of figures .......................................................................................................................... XI

List of abbreviations and symbols .......................................................................................... XIII

1 Introduction ......................................................................................................................1

1.1 Global CO2 production .................................................................................................1

1.2 Cimpor and cement production. ....................................................................................2

1.3 Carbon Dioxide ............................................................................................................8

2 State of art ..................................................................................................................... 11

2.1 Carbon Capture ......................................................................................................... 11

2.1.1 Post-combustion capture .................................................................................... 11

2.1.2 Oxy-combustion ................................................................................................. 12

2.1.3 Co-combustion and co-gasification ..................................................................... 12

2.2 Carbon Storage.......................................................................................................... 13

2.2.1 Storage in saline formations or aquifers .............................................................. 14

2.2.2 Injection into deep unminable seams or enhanced coal bed methane production 15

2.2.3 Use of CO2 in enhanced oil recovery ................................................................... 15

2.2.4 Depleted oil and gas reservoirs........................................................................... 17

2.3 Carbon Re-utilization .................................................................................................. 17

2.3.1 CO2 as inert........................................................................................................ 18

2.3.2 CO2 as working fluid ........................................................................................... 18

2.3.3 CO2 as Solvent ................................................................................................... 18

2.3.4 CO2 into more valuable products ........................................................................ 19

2.3.5 Convert CO2 into feedstock ................................................................................ 22

3 Tools ............................................................................................................................. 25

3.1 Aspen Plus® ............................................................................................................... 25

3.2 The Activated Economics ........................................................................................... 26

VIII

3.3 Aspen Energy Analyzer .............................................................................................. 27

3.4 Activated Exchanger Design & Rating ........................................................................ 27

4 Modeling ........................................................................................................................ 29

4.1 Process simulation ..................................................................................................... 29

4.2 Methanol Production Simulation ................................................................................. 31

4.2.1 The baseline scenario ........................................................................................ 31

4.3 The integration process .............................................................................................. 34

4.3.1 Scenario 2 .......................................................................................................... 34

4.3.2 Scenario 3 .......................................................................................................... 35

4.3.3 Scenario 4 .......................................................................................................... 36

4.3.4 Scenario 5 and 6 ................................................................................................ 36

4.3.5 Scenario 7 and 8 ................................................................................................ 37

4.4 Economic simulation .................................................................................................. 40

5 Results .......................................................................................................................... 49

6 Conclusion and Future work ........................................................................................... 56

Bibliography ............................................................................................................................ 57

Appendices ............................................................................................................................. 63

A1 - Scenario 1 ................................................................................................................... 63

A2 - Scenario 2 ................................................................................................................... 64

A3 - Scenario 3 ................................................................................................................... 65

A4 - Scenario 4 ................................................................................................................... 66

A5 - Scenario 5 ................................................................................................................... 67

A6 - Scenario 6 ................................................................................................................... 68

A7 - Scenario 7 ................................................................................................................... 69

A8 - Scenario 8 ................................................................................................................... 70

B1 - substitution of reboiler in heat exchanger. .................................................................... 71

C1 - B.E.U. TEMA sheet ...................................................................................................... 81

C2 - D.E.U. TEMA sheet ..................................................................................................... 82

IX

List of tables

Table 1: Composition of the flue gas feed ................................................................................ 29

Table 2: KPI's descriptions ...................................................................................................... 30

Table 3: Carbon dioxide conversion and selectivity's at 260°C and 330 bar ............................. 32

Table 4: KPI for the baseline simulation ................................................................................... 34

Table 5: reduced KPI in simulation 2 ....................................................................................... 35

Table 6: reduced KPI in simulation 7 ....................................................................................... 38

Table 7: KPI after the all the modifications, simulation 8 .......................................................... 40

Table 8: The KPI improvement between the baseline and scenario 9....................................... 49

Table 9: Summary of Cement Plant Costs With and Without CO2 Capture [62]........................ 54

Table 10: Economical evaluation of methanol plant scenario 9 ................................................ 54

Table 11: Economic improvement ........................................................................................... 55

XI

List of figures

Figure 1: Global CO2 production, IEA, Batelle 2002 [3]. .............................................................1

Figure 2: Investment in R&D [7]. ................................................................................................2

Figure 3: Cimpor cement manufacture process [10]. ..................................................................3

Figure 4: Typical rotary kiln [13]. ................................................................................................4

Figure 5: Decomposition of raw material components and formation of clinker phases with

increasing temperature [8]. ..................................................................................................................7

Figure 6: Global CO2 emissions from fossil fuel burning and cement production, giga tonnes

carbon per year [17] . ........................................................................................................................ 10

Figure 7: Schematic diagram of CO2 production, capture, and storage or re-utilization [2]. ....... 11

Figure 8: Overview of carbon capture technologies [24]. .......................................................... 12

Figure 9: IEA plan 2050 [25]. ................................................................................................... 13

Figure 10: CO2 storage overview [26]. ..................................................................................... 14

Figure 11: How CO2-EOR Works [32]. ..................................................................................... 16

Figure 12: Summary of CCU application typology [36]. ............................................................ 17

Figure 13: Carbon neutral cycle [38]. ....................................................................................... 19

Figure 14: Algae to biofuel process [46]. .................................................................................. 21

Figure 15: Inputs and Outputs of the Calera Process [48]. ....................................................... 22

Figure 16: CO2 to methanol simulation of [50]. ......................................................................... 25

Figure 17: TEMA Designation System Example [55]. ............................................................... 28

Figure 18: Methanol selectivity (SMeOH) at equilibrium (left) and Equilibrium conversion of CO2

(XCO2) (right) at different conditions of feed CO2:H2 ratio and pressure as a function of temperature

[40]. .................................................................................................................................................. 31

Figure 19: Stoichiometric reactor in Aspen Plus®. .................................................................... 32

Figure 20: Methanol reaction implementation in the stoichiometric reactor interface. ................ 33

Figure 21: Flowsheet compression unit [58]. ............................................................................ 34

Figure 22: Cooling water connection for hydrogen feed, simulation 3. ...................................... 35

Figure 23: The hydrogen recycle simulation 4. ......................................................................... 36

Figure 24: Super-heated steam integration simulation 5 and 6. ................................................ 37

Figure 25: H2 recycle on 50 bar simulation. .............................................................................. 38

Figure 26: Flowsheet Scenario 8. ............................................................................................ 39

Figure 27: Activated Economics simulation 8. .......................................................................... 41

Figure 28: Activated Exchanger Design &Rating of WW-gas-1. ............................................... 42

Figure 29: Costs and weights of WW-gas-1. ............................................................................ 42

Figure 30: Heat exchanger details by Aspen Energy Analyzer. ................................................ 44

Figure 31: Grid diagram in Aspen Energy Analyzer. ................................................................. 45

Figure 32: Hot and Cold composite curves of simulation 8. ...................................................... 46

Figure 33: Grand composite curve for Simulation 8. ................................................................. 46

XII

Figure 34: Hydrogen consumption over the 8 simulations. ....................................................... 49

Figure 35: Hydrogen efficiency. ............................................................................................... 50

Figure 36: Electrical power consumption over the 8 simulations. .............................................. 50

Figure 37: Thermal power consumption simulation 1 to 8. ....................................................... 51

Figure 38: Cooling water consumption simulation 1 to 9. ......................................................... 52

Figure 39: Heat exchangers size. ............................................................................................ 53

Figure 40: Cost tonne methanol ($/t). ...................................................................................... 55

Figure 41: Flowsheet scenario 1. ............................................................................................. 63



Figure 44: Flowsheet scenario 4. ............................................................................................ 66





Figure 49: Flowsheet for the substitution of the roboiler in destilation tower 1 in to a heat

exchanger. ........................................................................................................................................ 80

Figure 50: Design specification and sensivity test on the reboiler substitution. ......................... 80

Figure 51: B.E.U. TEMA sheet. ............................................................................................... 81

Figure 52: D.E.U. TEMA sheet ................................................................................................ 82

XIII

List of abbreviations and symbols

Abreviations Meaning

CCS Carbon Capture and Storage

CCU Carbon Capture and usage

CCUS Carbon Capture Usage and Storage

TEMA Tubular Exchanger Manufacturer Association

A.D. Anno Domini

IEAGHG IEA Greenhouse Gas R&D Programme

IEA International Energy Agency

GHG Greenhouse gas

CSI Cement Sustainability Initiative

UNIDO United Nations Industrial Development Organization

ECBM Enhanced coal bed methane production

EOR Enhanced oil recovery

CAP controlled atmosphere packaging

CFC chlorofluorocarbons

HCFC hydrochlorofluorocarbons

HFC hydrofluorocarbons

EGS Enhanced geothermal system

HDR Hot dry rock

LPG Liquefied Petroleum Gas

CRI Carbon Recycling International

DME Dimethyl ether

SCM Supplementary cementations material

STHE Shell-and-Tube Heat Exchangers

EDR Exchanger Design & Rating

KPI Key performance indicators

CW Cool water streams in simulations

WW Heat exchanger in simulations

NREL National Renewable Energy Laboratory

DOE Department of Energy

NASA National Aeronautics and Space Administration

R&D Research and Development

UV Ultraviolet light

XIV

Symbol Meaning

MMT Million metric tonne

Wt% Weight percentage

m meter

cm Centimeter

mm millimeter

kJ kilojoules

C° Degrees Celsius

g/mol Gram per mole

t/d Tonne per day

Mt/y Million tonnes per year

Mol% Mole percentage

km kilometer

Gt Giga tonne

V% Volume percentage

kW kilowatt

kWth Thermal power in kilowatt

kWel Electrical power in kilowatt

Mol/s Mole per second

kg/hr Kilogram per hour

∆H Reaction enthalpy

Gt Giga tonne

GtC/yr Giga tonnes carbon per year

SMeOH Methanol selectivity

XCO2 conversion of CO2

€M Million euros

€M/y Million euro per year

€t/y Euros per ton

GtC/yr Giga tonnes carbon per year

1

1 Introduction

1.1 Global CO2 production

It is well known around the world that the emissions of CO2 and other greenhouse gasses

support the global warming. The global warming is one of the most important issues of this century

and mostly the result of human activity. Our main source of energy is still the combustion of fossil

fuels, 83% [1], and will be for the next years to come. In addition to fossil fuel based heat and power

production processes, large stationary sources of CO2 emissions include natural gas sweetening,

hydrogen production for ammonia and ethylene oxide, oil refineries, iron and steel production facilities,

cement and limestone manufacturing plants [2].

Figure 1: Global CO2 production, IEA, Batelle 2002 [3].

Next to the global warming the CO2-emmions contribute urban smog, acid rain and health

problems as cancer. As a result, Carbon Capture and Storage (CCS), and Carbon Capture and

Utilization and storage (CCUS) are currently widely studied technologies to reduce anthropogenic

CO2-emissions [2].

As shown on the figure 1 the cement industry is currently one of the world’s main industrial

sources of carbon dioxide emissions. In 2011 the world production of cement got to 3.6 billion tonnes,

resulting in a carbon dioxide production of more than 2 billion tonnes from both the calcination of

limestone and the combustion of fuels [4] . Even if they have substantially reduced their emissions

over the years through alternative fuels or biofuels, clinker substitution and energy efficiency

improvement, there is still an increasing need to reduce them more. Although further reduction is

becoming limited, because CO2 has to be emitted during the decarbonation step in the clinker burning

2

process, CCS and CCU is becoming more and more important in this industry to reduce and prevent

CO2 emissions [5] [6].Many industry’s including the cement industry are facing potential large-scale

implementation of CO2 capture.

1.2 Cimpor and cement production.

Cimpor is Portugal’s largest manufacturer of cement and therefor a big producer of CO2. As the

whole world, Cimpor is concerned for the influence on the environment and the society. That’s why

Cimpor has already done a lot of investments to adapt their process in a more green process, (see

figure 2). This means that Cimpor increases the efficiency, specifically in terms of energy and is

focusing on the use of alternative fuels. It has made particular use of co-processing, which provides

both environmental and economic benefits [7].

Beside that Cimpor is also researching CCS, many of these technologies are not yet

commercially available nor is there any clear idea of their full potential on an industrial scale. One of

the examples in CCS research is in process in Portugal. In cooperation with another cement company

Cimpor is aiming to develop a pilot facility for the capture and storage of the CO2 emitted out of the

chimneys of the clinker kilns and reuse it in the production of bio-fuel and biomass from micro-algae

[3] [7].

Figure 2: Investment in R&D [7].

Cimpor's production process of cement in Portugal starts in the mining pits, see figure 3. Big

drills quarry out the raw materials. The raw materials are picked up by loading shovels and put into

dump-trucks or straight on a conveyor belt, like at the Cimpor plant in Alhandra. The raw material

consists of 75-79 weight % (wt%) limestone (which is rich of CaCO3) and 21-25 wt% clay (a source of

silica, alumina MgO and Fe2O3) [8]. After being quarried and transported to the main site, the raw

material gets crushed for the first time, followed by an initial blending stage. A natural mixture of the

two main materials occurs and is called marl. Quality reasons allow only a small deviation in the total

composition of the raw materials. As these still raw materials with some additives (sand, pyrite ash,

limestone of high lime content, etc.) passes by some hoppers and grinders, it gets simultaneously

3

dried and crushed until a very fine powder. This fine powder is than batched and stored in blending

and storage silos [9] [8] [10].

The clinker burning process needs a lot of thermal energy. For this energy many different fuels

are used in the cement industry. The most common are hard coal, lignite and petroleum cokes,

besides secondary fuels. They are delivered and stored in big coal silos before usage. These coals

and pet-cokes first need to be grinded by a grinder and filtered through bag-filters before they can be

injected and the ignition can follow in the kiln, providing an optimal thermal gradient for the clinker

process [11].

Figure 3: Cimpor cement manufacture process [10].

4

The next step is probably the most import step in de production process of cement, the clinker

burning process in the kiln. Here the fine powder gets a heat treatment, sintering, so that it converts

into clinker, where mineralogical small parts of cement can be found.

The raw mix enters at the upper end of the kiln and slowly works its way downwards to the

hottest area at the bottom over a period of 50-90 minutes, undergoing several different reactions as

the temperature increases. Typical dimensions for rotary kiln systems are: 50-100 m length and 3-7 m

in diameter, operating typically at a tilt of 1-3° from horizontal and with a rotation velocity of 2-4.5

revolutions per minute (rev/min). It is important that the mix moves slowly enough, to allow each

reaction to be completed at the right temperature. Because the initial reactions are endothermic, it is

difficult to heat the mix up to a higher temperature before a certain reaction is complete. The rotary

kiln thermal profile varies throughout its length, depending on the temperature and chemical reactions

involved during the process.

Where the wet process requires 5000 – 6000 kilojoule/kilogram (kJ/kg) of thermal energy per

kilogram clinker, the dry process has enough with 3400 – 5000 kJ/kg. This is because next to the

increased energy losses (waste gas and kiln wall) additional energy is needed for water evaporation

[12].

Figure 4: Typical rotary kiln [13].

The rotary kiln can be subdivided into several zones or regions that are exposed not only to

thermal and chemical wear but also to mechanical stresses (see Figure 4). The influence of one or

several of these factors, to minor or greater proportion determines the refractory lining type required

for each zone [9] [8] [14]:

• Decarbonation zone: from 300ºC to 1100°C (0 - 35 min)

5

This stage can occur either inside of the old wet process rotary kilns or in the preheater tower of

modern units. The stage consists of two steps: First step between 300°C and 650°C where the raw

meal heating occurs, accompanied by a dehydration reaction. The second step between 650°C and

1100°C, when the limestone decarbonation takes place generating CO2 and CaO [8] .

The first step is characterized by the following aspects:

- Presence of raw meal (there are no new mineral phases development);

- Erosion (due to raw meal flow at high velocities);

- low temperature;

- Evaporation and dehydration (of water) chemically bonded to the raw material.

In the second step, calcining, or calcination, is the process through which carbonates or other

compounds are decomposed by application of heat. Carbon dioxide is driven off from limestone

(CaCO3) and magnesium carbonate (MgCO3) , leaving free lime (CaO) and magnesia (MgO ) [9] .

Formation of CaO and CO2, see equation 1:

CaCO3ℎ𝑒𝑎𝑡→ 𝐶𝑎𝑂 + 𝐶𝑂2 (1)

At the temperature of 900°C the formation of 3CaO• Al2O3 starts.

Before entering the burning zone the complete kiln feed must have been calcined. This is

essential for the proper burning process and clinker formation. The successful calcining process,

which results in a grayish-green clinker, requires the appropriate temperature and an oxidizing

atmosphere for totally decomposition of the carbonates in the feed materials. Insufficient oxidizing

conditions yield a brown clinker that produces inferior cement [8] [14].

Burning zone

Burning process or often more used term sintering of the calcined kiln feed is actually a three-

stage process occurring in the hot end of the kiln (see figure 5). The area in where the burning occurs

can be divided into three sections: the upper transition zone, the sintering zone, and lower transition

zone. This zone results in the kiln feed becoming clinker. In the upper transition zone, interim-phase

formations occur while some calcination is still being completed. The upper transition zone is identified

by a quickly rise of temperature at the end of the calcining zone. The following stage of clinker

compound formation takes place in the sintering zone. In this zone the highest temperatures occur,

involving exothermic reactions. The last 3.05 to 6.1 m of the kiln represent the lower transition zone

and is characterized by a temperature drop, cooling step [9].

Upper transition zone: from 1100°C to 1300°C (35 - 40 min)

This is a very difficult and unstable zone. The temperature range varies from 1000°C to 1338°C,

but still problems of thermal overloads are frequent. These problems mainly are the result of the

6

specific design of the kiln main burner, the flame shape and to the fuel type that has been used.

Therefore, it is in this area where coating starts to develop as soon as first drops of liquid phase

appear [8].

Through the rise of the materials temperature, formation of secondary silicate phases occurs,

see equation 2:

2Ca + Si𝑂2ℎ𝑒𝑎𝑡→ 2𝐶𝑎𝑂 · 𝑆𝑖𝑂2

(2)

Sintering zone: from 1338ºC to 1450°C (40 - 50 min)

In this zone a full development of coating at 1450ºC (+) is expected. The liquid phase trace’s

represents the dissolution of 2CaO.SiO2, which indicates the reaction that generates clinker

3CaO.SiO2. The highest temperature in the kiln is reached at this area. Usually it should be around

1450ºC. Liquid phase is also around 25% at 1450ºC. If process is under control, coating will be stable

and able to protect the lining during the whole campaign. It is very important that there won’t be big

variability’s or uneven fuels types shifting, otherwise the coating will be unstable and refractories will

be submitted to an enormous thermo-chemical wear [9] [8].

Usually at this kiln zone it is possible to find:

- Presence of incipient liquid phase from 18 to 32%, free lime and C2S;

- Sintering and by the reaction of CaO and 2CaO.SiO2 within the melt to form ternary

silicates as alite, belite and tetracalcium aluminoferrates as shown in equation 3 and

equation 4:

2CaO. Si𝑂2 + 𝐶𝑎𝑂ℎ𝑒𝑎𝑡+𝑡𝑖𝑚𝑒→ 3𝐶𝑎𝑂 · 𝑆𝑖𝑂2 (3)

3CaO. Al2𝑂3 + 𝐶𝑎𝑂 + 𝐹𝑒2𝑂3ℎ𝑒𝑎𝑡+𝑡𝑖𝑚𝑒→ 4𝐶𝑎𝑂. 𝐴𝑙2𝑂3. 𝐹𝑒2𝑂3 (4)

- Clinker liquid phase infiltration and coating formation;

- Chemical attacks by alkaline sulfates;

- High operational temperature.

Lower transition zone from: 1450°C to 1250°C (50 - 60 min)

This area usually operates between 1450°C and 1250°C. Crystallization of the clinker start

around 1250 °C. Even though some spores of the liquid phase can still occur. At this stage the

chemical activity is quiet low, considering the biggest part of 3CaO.SiO2 has already been formed with

a remaining amount of free lime around of 1%. Still it is a zone with remarkable temperature variations

as a consequence of the secondary air temperature coming from the cooler [9] [14].

7

This area is characterized by the following aspects:

- Presence of the clinker liquid phase;

- Chemical attacks by alkaline sulfates;

- Frequent temperature variations when flame impinges over the brick;

- Continuous thermal shock;

- Mechanical stress imposed by the tire station and kiln shell ovality.

• Pre-cooling zone from: 1250 °C to 1000°C

Cooling of the clinker commences in the precooling zone of the kiln. The length of this zone and

the cooling rate depend on several factors, for example, on the primary burner position, the type of

clinker cooler and the secondary air temperature entering into the kiln. The clinker leaves the kiln at a

temperature of about 1350°C (sintering zone closes at the kiln exit) to 1150°C. With old grate coolers it

was around 700°C. The main purpose of this zone is to finalize the crystallization of the various

mineral phases formed in the kiln [8].

The main characteristics of this kiln zone are:

- High abrasion and erosion by the clinker dust carried by secondary and tertiary airs

- Frequent thermal shocks;

- High mechanical stresses (compression/traction).

Most of the modern furnaces are equipped with high efficiency coolers, these coolers are not

inside the rotary kiln but in the first cooling compartment.

Figure 5: Decomposition of raw material components and formation of clinker phases with increasing temperature [8].

8

After precooling, the clinker is further cooled in rotary, planetary, or grate-type coolers by forcing

air through the unit by dedicated cooler fans. The cooler is designed to drop the temperature of the

clinker from 1000 °C to 150 °C. The cooling process conditions can significantly influence the quality

of the clinker. Generally, faster cooling rates result in a higher quality clinker [8].

At this point in the process the materials have been formed into all the required minerals to

make cement. The whole process of decomposition of raw material components and formation of

clinker phases is illustrated in figure 5. Like cement, the clinker will react with water and harden, but

because it is composed of 1-3 cm diameter fragments it is too coarse to be used, so subsequently the

clinker is then transferred by a conveyor to the finish mill [11].

To produce the final product it is necessary to grind the clinker for approximately 30 minutes in

large tube mills. The result is a very fine powder, 94 to 98 percent particles with diameters less than

0.044 mm. Together with the grinding process the clinker is mixed with a binding regulator, about five

percent gypsum (CaSO4.2H2O) and perhaps other additives (limestone filler, fly-ash, steel-mill slag,

etc,) gives rise to the various types of cement that meet the various standards. The gypsum is added

to retard the setting time of the cement, thereby making it more suitable for common construction

applications. The cement grinding process is highly energy intensive. Therefor water is added to both

the inlet and outlet ends of the mill to cool the product and the mill itself [10].

The cement is then pumped to storage silos and it may then be sold in bulk or packed in paper

bags that are shipped out on pallets or packet-packed. Dispatch can be by lorry, train or ship, as

necessary [10]

1.3 Carbon Dioxide

The existence of carbon dioxide has been known since primitive times. In the first century A.D,

Pliny the Elder wrote of lethal vapors that were coming out of caverns. Many centuries later, van

Helmont (1577 – 1644) produced this gas by various pathways and proved it was the same gas as the

one that is issued from caverns and mines. Josphen Black discovered in 1757 that carbon dioxide is

produced by animal respiration and microbial fermentation. He mentioned the lethal effect of the gas

on animal life and called it “fixed air”. The first time that carbon dioxide was used for “fixed air”, was by

Antoine Lavoisier, by showing that it was produced when carbon is heated in oxygen. The occurring

gas, which he named carbonic acid, was found to contain 23.5 – 28.9 parts by mass carbon and 71.1

– 76.5 parts by mass oxygen. The first really introduction in the industry started with the experiments

of Faraday on the liquefaction of gases. He realized in liquefying CO2 in a bent glass tube. Further

experiences were done by Thilorier, he could convert the gaseous carbon dioxide into solid phase.

The solid carbon dioxide appeared as a white, flocculant, easily compressible mass, better known as

dry ice. With the invention of the first mechanical compressor by Johann Natterer the interest into the

liquid uses of Carbon dioxide increased. So in 1884 the first factory for the production of liquid carbon

dioxide was established by Raydt in Germany. As more and more uses were found, the development

of the compressors and the industry continued. Since the turn of the century, many uses for carbon

9

dioxide have been identified, and several other methods of manufacture have assumed commercial

importance [15].

Carbon dioxide, CO2, molar mass: 44.010 grams/mole (g/mol), is gaseous phase at normal

temperature and is not very reactive, inert. CO2 is a colorless, odorless, nonflammable gas with a

slightly sour taste. The molecule is relatively stable and won’t break down easily in smaller compounds

[15]. Only by heating up, ultraviolet light, or electrical discharge the carbon dioxide might break down

as equation 5:

𝐶𝑂2 ↽⃑⃑⃑ 𝐶𝑂 + 0.5𝑂2 (5)

High temperatures or the use of catalysts can introduce also other reactions with carbon

dioxide. The reverse of the water-gas shift reactions, where carbon dioxide reacts with hydrogen

results in carbon monoxide, shown in equation 6.

𝐶𝑂2 +𝐻2 ↽⃑⃑⃑ 𝐶𝑂 + 𝐻2𝑂 (6)

Or at an elevated temperature carbon monoxide is formed by a reaction with carbon, equation

7.

𝐶𝑂2 + 𝐶 ↽⃑⃑⃑ 2𝐶𝑂 (7)

Carbon dioxide is also very important in the biosphere, so it is a vital link in the life cycle of

plants and animals and, as such, carbon dioxide is basic to all life on earth. The leaves of the plants

absorb the carbon dioxide out of the air. Subsequently the chlorophyll in the leaves in combination

with the energy of the sunlight, ultraviolet light (uv), converts the absorbed carbon dioxide in a reaction

with water into glucose, photosynthesis, see equation 8.

6𝐶𝑂2 + 6𝐻2𝑂𝑐ℎ𝑙𝑜𝑟𝑜𝑝ℎ𝑦𝑙𝑙 𝑢𝑣 → 𝐶6𝐻12𝑂6 + 6𝑂2 (8)

∆H=2803 kJ/mole

The glucose is used by plants to produce sugar and starch, key elements of the life cycle.

These elements which are in plants will be eaten by animals. Animals, in turn, will break the sugars

and starch back to carbon dioxide. Thus, carbon is being continually exchanged between the bodies of

plants and animals. Atmospheric carbon dioxide is the medium of this exchange [15].

The problem is that there are plenty of other processes nowadays in the world that as well

produce carbon dioxide, which results in an excess of carbon dioxide in the atmosphere. The principal

cause of this excess, is the combustion of fossil fuels. Since the industrial revolution around 1750, the

natural coal, gas and fuel has increased dramatically [16]. Measurements of the last 25 years show us

that there used to be an annual increase on the order of 1 % of carbon dioxide emissions a year but in

the last 10 years it grew to 2% and more the years after, see figure 6 Global CO2 emissions from fossil

fuel burning and cement production, giga tonnes carbon per year (GtC/yr) [17].

10

Figure 6: Global CO2 emissions from fossil fuel burning and cement production, giga tonnes carbon per year [17] .

The problem is that the overload of carbon dioxide in the atmosphere promotes the commonly

known “greenhouse effect”. When sunlight reaches earth’s surface, it can either be reflected back into

space or absorbed by the Earth. Carbon dioxide is transparent to the sun’s incoming short wave

radiation, but once absorbed, the planet releases some of the energy back into the atmosphere as

heat, also called infrared radiation or long wave radiation. Greenhouse gases like water vapor (H2O),

carbon dioxide, and methane (CH4) absorb the heat energy, which results in slowing or preventing the

loss of heat to space. In this way, greenhouse gases act like a blanket or like the glass in a

greenhouse, making Earth warmer than it would otherwise be and result in climate changes [18].

The changing climate effects society and ecosystems in a broad variety of ways. For example

climate change heated up the average global temperature which subsequently let the polar capes

melt. There for the coasts will likely experience stronger hurricanes and sea level rise. The changing

climate also induce the increase or decrease of rainfall, influence agricultural crop yields, affect human

health, cause changes to forests and other ecosystems, or even impact our energy supply. Climate-

related impacts are occurring across regions of the country and across many sectors of our economy.

Many state and local governments are already preparing for the impacts of climate change through

"adaptation," which is planning for the changes that are expected to occur but not solving the core

problem [19].

Nowadays several scientists do believe that some actions should be carried out to find a way in

reducing the global emissions of greenhouse gases and therefore minimize the climate change. This

can be done in many different ways. Carbon capture, storage and reuse are important points on the

agenda of plenty of scientist and companies. This is proved by the IEA Greenhouse Gas R&D

Program (IEAGHG), an international collaborative research program established in 1991 as an

implementing Agreement under the International Energy Agency (IEA). The role of the Program is to

evaluate technologies that can reduce greenhouse gas emissions derived from the use of fossil fuels

[20].

11

2 State of art

2.1 Carbon Capture

Many industries including the cement industry are facing potential large-scale implementation of

CO2 capture. Figure 7 presents the simplified pathways of CO2 production, capture or separation, and

storage or re-utilization [2].

Figure 7: Schematic diagram of CO2 production, capture, and storage or re-utilization [2].

Considering the three most advanced capture technologies (illustrated in figure 8), post-

combustion and oxy-combustion capture processes seem to be useful to the cement industry. Pre-

combustion techniques are inadequate because CO2 is produced during the conversion of limestone

to calcium oxide in the clinker burning process [2].

2.1.1 Post-combustion capture

Post combustion is a downstream process (chemical absorption, adsorption, membranes, etc.),

which won’t affect the overall cement production process. Therefore this technology is applicable not

only for new kilns but also for plants that already exist [21].

The CO2 concentration of the cement plant flue gasses, around 25 mole percent (mol%), are

quiet high compared to the concentration of flue gasses in a coal fired power plant, about 14 mol%.

Still the same process can be used in the cement industries, with a monoethanolanie solvent. But this

amine scrubbing process is widely used for small installation with a production of 400t/d of CO2. For a

production of 1 million tonnes a year (Mt/y) of cement, there is a scale-up needed around 3000 tonnes

a day (t/d) of CO2 [21].

The post combustion capture process has a really pure CO2 stream, generally 99.9% dry basis.

In general 85% of CO2 can be captured in power plants, but from studies of IEA GHG’s on post

12

combustion and other published work it seems that 95% capture is possible without any serious

investments. It is known that the CO2 concentration in cement industry flue gas is higher than the one

in power plants flue gas, therefore it can be assumed that the amount of CO2 captured is equal or

maybe even higher than in power plants [21].

2.1.2 Oxy-combustion

The second option is oxyfuel technology or oxy-combustion. With this technology the CO2 will

be captured at the cement kilns. The pure CO2 is achieved by using pure oxygen, that is been

provided by an air separation unit, instead of air. Recycling a certain amount of CO2 rich flue gases will

have a huge impact on the clinker burning process, mainly thermal energy efficiency. During the last

years the oxyfuel technology has been investigated in the power plants. Which means that it results

can be implemented in the cement industry, kilns. Still there is a lot of research needed before the

scale up to industrial industry is possible [22]. Mainly due its advantages of improving the energy

demand, the oxyfuel technology compared to other capture methods is seen as a promising method

for the long term perspective [21].

2.1.3 Co-combustion and co-gasification

Besides using the different capture methods separately, it is also possible to combine them. The

so called hybrid solutions may lead to higher energy efficiency. One of the possibilities that would be

possible is the combination of moderate oxygen enriched combustion with or without flue gas recycle

followed by a post combustion capture. This could lead to a higher CO2 concentration at lower volume

flows in the flue gas. The use of pure oxygen will reduce the amount of nitrogen in the flue gas. This

would reduce the size and energy consumption of the post combustion capture, like scrubbers. Also

would it enable alternative ways for post combustion capture technologies such as the use of

membranes. It was found that in the power sector the combination of oxygen enrichment and the use

of a membrane for post combustion capture would be more efficient. However this kind of research

has not yet been done in the cement sector [23].

Figure 8: Overview of carbon capture technologies [24].

13

The IEA and the Cement Sustainability Initiative (CSI) member companies have cooperated to

develop a plan for the cement industry for carbon emissions reductions up to 2050 [25]. According to

this plan, the cement industries will have to invest time and money in the years to come, so they can

reduce the CO2 emissions to the level of the agreement. The major part of the reduction efforts is seen

as being provided by the application of CCS technologies (see Figure 9). In 2050, about 50% of all

cement kilns in Europe, North America and Australia have to be equipped with carbon capture

technologies. Other countries with a big cement industry like china and India will have CCS in 20 % of

the kiln lines [25].

In 2011 there was a new study on CCS in industrial applications by the United Nations Industrial

Development Organization (UNIDO) in cooperation with the IEA. According this study the cement

industry is seen to have the potential to capture 500 million tons a year of CO2 in 2050 by the

application of CCS technologies. [25]

Figure 9: IEA plan 2050 [25].

Although there are a lot of possibilities where CCS can be implemented in the cement industry it

is obvious even today that the estimated costs for CCS application in the cement industry will be

extremely high. From today’s view, cement production costs would roughly double [25].

2.2 Carbon Storage

After the capture, transporting carbon dioxide is a vital step in the CCS process. In CCS

transport of CO2 in pipelines is a well-known and mature technology, with significant experience from

more than 6 000 kilometers (km) of CO2 pipes in the United States [26]. There is also experience,

although limited, with transport of CO2 using offshore pipelines in the Snøhvit project in Norway [27].

CO2 is also transported by ship, but in small quantities. Although the technical requirements and

conditions for CO2 transport by ship has improved recently [26]. For further improvement in CCS, it will

be necessary to link CO2 pipeline networks across national borders and reduce the costs of shipping

14

transportation infrastructure, by temporary storage and liquefaction facilities. Here for the main goal

will be to construct long term international strategies that will link the captured carbon dioxide to the

places where it can be stored. Government-led, national or regional planning exercises are required in

this regard. [26]

The storage and re-use of the carbon dioxide has found lots of interest the last decennia,

supported by the increasing climate problems. By the research that has been done there are already

applications with potential of storing and reusing the carbon dioxide. The storage of carbon dioxide

happens deep under the ground in geological reservoirs and it will stay there for millennia. It has been

proven that the carbon dioxide will rise back to the atmosphere. Migration to the surface would take

millions of years [26]. The main goal of the storage of the carbon dioxide is to achieve the 2DS

objectives. The 2DS objective is an internationally agreed target of limiting the global temperature rise

to 2°C above preindustrial levels by 2025 or 20°C by 2050. [28]

The different possibilities of carbon storage are illustrated in figure 10.

Figure 10: CO2 storage overview [26].

2.2.1 Storage in saline formations or aquifers

Deep saline aquifers offer the largest storage potential of all the geological CO2 storage options

and are widely distributed throughout the globe in all sedimentary basins. Injection of CO2 into deep

saline aquifers involves CO2 as a supercritical fluid that is less dense and less viscous than the

resident formation water. Thus, even if the injected CO2 itself is safely trapped in suitable geological

structures, pressure changes and brine displacement may affect shallow groundwater resources, for

example, by increasing the rate of discharge into a lake or stream, or by mixing of brine into drinking

water aquifers [29].

15

2.2.2 Injection into deep unminable seams or enhanced coal bed methane production

Storing CO2 in unminable coal seams is another option for the storage of carbon dioxide. These

coals seams often contain gasses, such as methane. The gasses are held in pores on the surface of

coal and in small fractures in the seams. Therefore, if CO2 is injected in the coal seams the gasses will

be displaced. For example the methane can be recovered as a free gas, this process, is well known as

CO2-enhanced coal bed methane production (CO2-ECBM). 90% or more of the methane in the coal

seam can be recovered with this technique. The methane can be used for thermal heating, power

generation or sold to cover the cost of CO2-injecting [30].

2.2.3 Use of CO2 in enhanced oil recovery

Off the projects that are under construction or at an advanced stage of planning, 70% (16 of 22)

intend to use captured CO2 to improve recovery of oil in mature fields better called as enhanced oil

recovery (EOR). CO2-EOR projects merit cautious treatment as an indicator of progress in CCS

deployment [31].

For the production and development of crude oil in the United States oil reservoirs, the United

States has up to three distinct phases: primary, secondary, and tertiary (or enhanced) recovery. The

primary state is responsible for about 10 percent of reservoir's original oil, by the natural pressure of

the reservoir or gravity the oil will flow into a wellbore. Artificial lift techniques, pumps for example, will

bring the oil to the surface. Secondary recovery techniques include a more productive way of

recovering, as it can produce around 40 percent of the original oil in place. In this manner, water or

gas is injected to displace oil and drive it to a production wellbore [31].

After exploiting the much easier-to-produce oil from U.S. oil fields, producers have fiend a third

way for producing 30 to 60 percent, or more, of the reservoir's original oil in place. Enhanced oil

recovery, techniques that offer prospects, three major categories of EOR have been found to be

commercially successful to varying degrees:

Thermal recovery involves the injection of steam to lower the viscosity of the sticky oil, and

makes it easier to let the oil flow through the reservoir. Thermal techniques account for over 40

percent of U.S. EOR production, primarily in California [32].

The other 60 percent of EOR production in the United States is gas injection, which uses

gasses such as natural gas, nitrogen, or carbon dioxide CO2. These gases will expand in the

reservoirs to push additional oil to a production wellbore. Gases can also be injected to lower the

viscosity of the oil and improves its flow rate. [32]

16

Figure 11: How CO2-EOR Works [32].

In the United States EOR production has also chemical route. Either way long-chained

molecules, polymers, are injected to increase the effectiveness of water floods, or detergent-like

surfactants are used to help lower the surface tension that often prevents oil droplets from moving

through a reservoir. Each of these techniques has been hampered by its relatively high cost and, in

some cases, by the unpredictability of its effectiveness [32].

The EOR technique that is attracting the most new market interest is CO2-EOR, shown in figure

11. In the U.S., there are about 114 active commercial CO2 injection projects that together inject over

2 billion cubic feet of CO2 and produce over 280,000 barrels of oil per day [33]. First tried in 1972 in

Scurry County, Texas, CO2 injection has been used successfully throughout the Permian Basin of

West Texas and eastern New Mexico, and is now being pursued to a limited extent in Kansas,

Mississippi, Wyoming, Oklahoma, Colorado, Utah, Montana, Alaska, and Pennsylvania [32].

In the EOR process 51% of the CO2 that was injected is recovered and recycled in the process.

The recovered CO2 appears in the produced oil, water and gas flows. The remaining part of the CO2

has stored itself in the following way, [32]:

- 12% in oil phase

- 18% in brine

- 70% in gaseous phase

- Mineral carbonation <1% of that stored

In 2009 IEAGHG study gave a global capacity of 140GTCO2 storage in depleted oil fields [32].

17

2.2.4 Depleted oil and gas reservoirs

Until recently, most of the CO2 used for EOR has come from naturally-occurring reservoirs. But

new technologies are being developed to produce CO2 from industrial applications such as natural gas

processing, fertilizer, ethanol, and hydrogen plants in locations where naturally occurring reservoirs

are not available. One demonstration at the Dakota Gasification Company's plant in Beulah, North

Dakota is producing CO2 and delivering it by a 204-mile pipeline to the Weyburn oil field in

Saskatchewan, Canada. Encana, the field's operator, is injecting the CO2 to extend the field's

productive life, hoping to add another 25 years and as much as 130 million barrels of oil that might

otherwise have been abandoned [34]. The U.S. National Energy Technology Laboratory estimates

that for every metric ton of CO2, 1-2 barrels can be produced at a cost of $8.20 / barrel [35].

EOR is generally used in depleted oil reservoirs. But CO2 can as well be injected in depleted

gas reservoirs. Globally, 900 giga tonne (Gt) of CO2 could be stored in depleted natural gas fields,

substantially more than in depleted oil fields. There are investigations going on with regard on the

possibility to inject CO2 into gas field that are almost empty for enhanced gas production, without

contaminating the residual gas. Still further research is needed to see if the enhanced gas production

by CO2 injection is economically similar to EOR [34].

2.3 Carbon Re-utilization

Beside the storage, the re-use of carbon dioxide gets more and more attention. In general these

are the most used applications of CO2, see figure12. Additionally, CCU is widely viewed as a suite of

technologies that can provide support for CCS.

Figure 12: Summary of CCU application typology [36].

18

The applications of Carbon dioxide can, as showed on figure 12, be divided in to two main

groups. Or the carbon molecule gets converted into a new more valuable product or the molecules

specific properties will be used [36].

2.3.1 CO2 as inert

One important characteristic of CO2 is that it is inert and so harmless for humans. Therefore

CO2 often is used in gaseous phase in the food industry, more specific the packaging industry. Carbon

dioxide is commonly used in modified atmosphere packaging (MAP) and in controlled atmosphere

packaging (CAP) because of its ability to inhibit growth of bacteria that cause spoilage. In the wine

industry CO2 is a perfect seal gas to prevent oxidation of the wine, during its maturation. Gaseous

CO2 is an important alternative in replacing environmentally harmful refrigerants such as

chlorofluorocarbons (CFC), hydrochlorofluorocarbons (HCFC) and hydrofluorocarbons (HFC). The

development has shown potential in several applications both in terms of system efficiency and

system cost, thus making it a viable alternative for applications as a refrigerant gas for large industrial

air conditioning and refrigeration systems. [37]

2.3.2 CO2 as working fluid

CO2 in liquid phase has an important utilization in the power generation industry. This power

generation process is an enhanced geothermal system (EGS) using CO2 instead of water or brine as

working fluid, which is a variant of “hot rock” geothermal energy systems. Supercritical CO2 is

circulated as the heat exchange fluid to recover the geothermal heat from the reservoir. The

geothermal energy will be converted into electricity in a power plant at the surface. EGSCO2 has

benefits ahead of water-based Hot Dry Rock (HDR) system, because the use of CO2 reduces the

pumping requirements compared with water-based systems. These benefits of CO2-use will increase

the generation efficiency at the surface and reduce the heat loss in the geothermal reservoir. Another

benefit of EGSCO2 is that with CO2 much higher flow rates can be achieved than with water due the

lower viscosity of CO2. For similar reasons, there is growing interest in the use of supercritical CO2 in

closed loop power cycles as a replacement for steam, resulting less energy needed from fossil fuel-

fired or nuclear power plants [36].

2.3.3 CO2 as Solvent

It is possible to use CO2 as a solvent, for example for the production of Enhanced hydrocarbon.

This group of technologies will use CO2 as solvent to increase recovery of hydrocarbons from the

subsurface, as the already mentioned CO2-EOR and its variants, which are a combination of storage

and re-use.

19

2.3.4 CO2 into more valuable products

As already mentioned the most CO2-emissions are derived from the combustion of fossil fuels.

The conversions of CO2 back into more valuable products therefore will most of the time need

hydrogen. The production of hydrogen is an intensive energetic process. This energy has all kind of

different sources, but in general it comes from natural gas combustion, with the consequence of new

CO2-emissions. To avoid more CO2-emissions, the electrolysis of water (hydrolysis) is promising

solutions. The power for this electrolysis will have to be sun energy or renewable energy. Figure 13

gives good overview of the carbon neutral cycle. In this way the following products can be made and

used in widely spread applications [38].

Figure 13: Carbon neutral cycle [38].

2.3.4.1 Methanol

A popular liquid fuel produced in the CO2 re-use industry is methanol. This product is a starting

molecule for a diversity of hydrocarbons and a second generation fuel that with kind of similar

properties as Liquefied Petroleum Gas (LPG). There exists an productive process for the synthesis of

methanol by the route of a continuous catalytic hydrogenation of CO2 under high-pressure conditions,

up to 360 bar, over co-precipitated Cu/ZnO/Al2O3 catalysts. The catalytic system is an efficient and

highly productive processing strategy for CO2 conversion at a single step. The reaction is exothermic

with an enthalpy of -130.7 kJ/mol, equation 9 [39]:

20

𝐶𝑂2 + 3𝐻2 𝐶𝑢/𝑍𝑛𝑂/𝐴𝑙2𝑂3→

𝐶𝐻3𝑂𝐻 +𝐻2𝑂 (9)

∆H= -130.7 kJ/mol

Under these conditions the one-pass CO2 conversion is bigger than 95% and methanol

selectivity can get higher than 98%. Due the possibility of a reverse water-gas shift (equation 10),

producing water and carbon monoxide from CO2, there needs to be an excess of hydrogen to have

this conversion to methanol. The excess of hydrogen, H2/CO2 = 10/ [40], results into a second reaction

in the reactor, producing methanol in the following way from carbon monoxide, equation 10:

𝐶𝑂 + 2𝐻2 𝐶𝑢/𝑍𝑛𝑂/𝐴𝑙2𝑂3→

𝐶𝐻3𝑂𝐻 (10)

∆H= -92 kJ/mole

The biggest issue is to have a renewable source of energy, which is not always available, for

example in Germany with its sun or wind energy. Here for Carbon Recycling International (CRI) has

plans with industrial partners from Germany, Spain and Belgium to implement its Emissions-to-Liquids

technology for recycling carbon-dioxide emissions from a coal-fired power plant in Germany. Carbon

Recycling International is known as a producer of renewable methanol, under the brand name

Vulcanol. CRI’s Emissions-to-Liquids production facility is located at the geothermal power station in

the town of Grindavik in Reykjanes, Iceland. The plant was constructed in 2012 and was able to

produce around 2 million liters of renewable methanol. CRI plans to expand the plant to produce more

than 5 million liters a year by 2014 and to reclaim 5.5 thousand tonnes of carbon dioxide a year from

the atmosphere [41].

2.3.4.2 Dimetheyl ether

The Synthetic gasoline dimethyl ether (DME), may be produced from coal-derived flue gas via

the methanol-to-gasoline (MTG) process of ExxonMobil, a two-step process that in a first step

converts CO2 into methanol( equation 9) followed by second step that produces DME (equation 11)

[42]. DME is the simplest ether and is considered a leading alternative to petroleum-based fuels and

liquefied natural gas. Its physical properties are similar to LPG and DME is considered as a substitute

for diesel fuel because it has a cetane number between 55 and 60, although some modifications of the

engine are required. The methanol that is mentioned her above can so be dehydrated, the reaction is

carried out catalytically over varied solid acids such as alumina or phosphoric acids, γ-Al2O3 [43].

2𝐶𝐻3𝑂𝐻 γ−Al2O3→ 𝐶𝐻3𝑂𝐶𝐻3 +𝐻2𝑂 (11)

∆H= -23.5 kJ/mole

In Canada Blue Fuel Energy plans to produce low-carbon methanol (Blue Fuel methanol) and

low-carbon DME (Blue Fuel DME), using renewable electricity from wind and waste carbon dioxide.

The BioDME project in Sweden demonstrates the use of renewable DME (by biomass power) as an

ultra-clean transportation fuel [44].

21

2.3.4.3 Methane

Methane can be produced from carbon dioxide and hydrogen, by the Sabatier’s reaction,

equation 12. Here again, the same problem occurs that the production of hydrogen is very expensive,

so renewable energy will be necessary for the electrolysis.

𝐶𝑂2 + 4𝐻2 𝑁𝑖−𝑆𝑖𝑂2→

𝐶𝐻4 + 2𝐻2𝑂 (12)

∆H= -164.9 kJ/mole

Several other catalysts have been studied: Ni-La2O3-Ru, Cu-Zn-Cr, Fe-Cu

The most valuable use of methane is to convert it into energy, making methane a biogas. A

biogas has a wide range of application, as for example a heat provider. Methane as fuel gas for cars is

another useful application. The world’s largest power-to-gas plant currently in operation is situated in

Germany, Werlte. Audi produces fuel gas for cars, e-gas, produced with renewable energy, mainly

wind energy. The Audi e-gas plant produces about 1,000 metric tons of e-gas per year, chemically

binding some 2,800 metric tons of CO2. This roughly corresponds to the amount that a forest of over

220,000 beech trees absorbs in one year. Water and oxygen are the only by-products [45].

2.3.4.4 Use of Algae

Algae have been widely studied in recent years as a source of biomass for the production of

biofuels. The algae production process for biofuels and biomass is based on a cultivation of algae in

typically saline or brackish water in open ponds or closed bioreactors, where CO2 is bubbled through

to accelerate biomass production. The ponds or bioreactors have the highest yield in saturated carbon

dioxide conditions. The lipid fraction of the biomass can be used to make biodiesel and other liquid

fuel substitutes. Another part of the cultivation of these photosynthetic microorganisms, as algae are

using only solar energy and CO2, is the directly excrete of hydrocarbons that can be used as fossil fuel

substitutes.

Figure 14: Algae to biofuel process [46].

After cultivating the in open ponds, Algae are separated from a growth media, lipids are

extracted, and transesterified to produce biodiesel (shown on figure 14). The non-lipid residual

biomass is processed into usable co-products such as biogas or animal feed [46].

22

For example Seambiotic grows microalgal cultures in open ponds using flue gases like carbon

dioxide and nitrogen from a nearby coal plant as feedstock. Its 1000-square-meter facility produces

roughly 23,000 grams of algae per day. Three tons of algal biomass would yield around 100 to 200

gallons of biofuel. It recently formed a partnership with National Aeronautics and Space Administration

(NASA) to optimize the growth rates of its microalgae. The company is currently in transition from the

pilot plant stage to commercial scale algae cultivation and production [47].

2.3.5 Convert CO2 into feedstock

A final way of CCU that delivers benefits which leads to reductions in GHG-emissions is the

conversion of carbon dioxide in more intensive forms of production of intermediates within a value

chain, for example bulk chemicals or polymers [36].

2.3.5.1 CO2 mineralization

Mineralization is a technology that relies on the accelerated chemical weathering of certain

minerals using CO2. CO2 mineralization is a relatively new technology in the field of CO2 geological

sequestration. Compared with other CCS technologies, it has the advantage of safely storing CO2 for

a very long time, if not infinite, by converting CO2 into a solid phase. Through a reaction between rich

calcium and magnesium ions in natural alkaline ores and alkaline wastes, CO2 could be converted into

stable solid carbonates, such as magnesium carbonate and calcium carbonate, shown on figure15.

These are minerals that can be reused in the cement production process as alternative feedstock [48].

Figure 15: Inputs and Outputs of the Calera Process [48].

23

The company Calera Corporation based in Los Gatos, California, operates a pilot plant next to a

1000 MW power plant at nearby Moss Landing that can operate continuously and currently can

produce an average of five ton of supplementary cementations material (SCM) per day from the

carbon dioxide emissions. At the pilot level CO2 capture rates were achieved of greater than 50% and

absorption rates above 90%. The conversion of CO2 to calcium carbonate requires a source of

alkalinity and calcium. One option is the use of industrial waste streams that contain both alkalinity and

calcium, for example in the form of calcium hydroxide or calcium chloride (can be naturally occurring

or can also be found in the waste streams of existing chemical processes). In prospect the Portland

cement consumption will nearly be 200 million metric tons by 2020, a 20% market penetration with a

SCM replacement level of 50% would reduce CO2 emissions by 30 million metric tons [48].

25

3 Tools

3.1 Aspen Plus®

For engineering calculations and simulations the Aspen Tech software is the basis for designing

new processes or upgrading existing processes to improve their performance. The program Aspen

Plus is the market-leading chemical process optimization software used by the bulk, fine, specialty, &

biochemical industries, as well as the polymers industry for the design, operation, and optimization of

safe, profitable manufacturing facilities. The software in the aggregate is used for constructing models

and making business, using the results of simulations [49].

All models of aspen software are based on the knowledge of technological processes, and

combines all engineering innovations and advances in information technology, and yield reliable

results, tested in real industrial plants.

The simulations in this research are based on a CO2 to methanol simulation out of literature

[50], figure 16, and made in Aspen plus®. Starting from this simulation, several adjustments are

introduced.

Figure 16: CO2 to methanol simulation of [50].

The main units that have been used in this simulation are reactor, gas compressors, heat

exchangers, separators, distillation towers, valves and mixers. For each model its specific inputs are

given, like conversion, mass flows, reflux rates, etc.

26

3.2 The Activated Economics

This program has been introduced in the research project to have an idea of the equipment

costs of the CO2 to methanol process.

With Activated Economics, it is possible to run simultaneous process cost evaluation while

building a model in Aspen HYSYS or Aspen Plus. Activated Economic Analysis runs concurrently with

the user simulation, updating whenever the model is changed. This allows the user to immediately

view the economic impact of design changes for faster and more efficient process modeling and

process designs optimized for capital and operating expenditure. Activated Economics gives process

engineers insight into the capital and operating costs associated with their models throughout the

design process, keeping them on track with the most economically feasible designs before they send

equipment lists and process designs to the estimating department. By using Activated Economics

during design, process engineers gain the advantage of pre-screening their own models before

sending them off to other departments for more thorough cost estimates and analysis. This integrated

workflow increases engineering efficiency by allowing process engineers to quickly screen out

infeasible designs based on costs or constructability of equipment. Ultimately, the workflow of iterating

through models based on feedback from other departments becomes streamlined, allowing everyone

on the team (from process engineers to equipment and estimating specialists) to spend more time

working on higher-quality design cases [51].

By using this tool, several parameters can be measured and afterwards compared between the

various simulations that have been run. The main parameters that will be examined are:

- The total capital cost: this number implements the total cost of all the equipment that is

needed in the process, reactor, distillation tower, compressors, etc.

- The total operating cost: this number represents the purchase of all the raw materials,

utilities of the equipment’s, the labor and maintenance costs.

- The total raw materials cost: the amount of the operating cost that needs to be purchased.

- The total product sales: the total amount of product that can be sold on the marked. This

includes mainly methanol, but oxygen or steam can be sold as well.

- The total utilities cost: It represents the amount of the operating cost, which are spent on

utilities like steam, cooling water, electricity.

Obtaining all these results and calculating the process parameters can lead the research to

several optimizations in different areas of the plant. Microsoft Excel is applied to insert the results and

process parameters into clear and organized graphics and tables.

27

3.3 Aspen Energy Analyzer

Next to the cost of the equipment, the energy optimization is also an important research subject.

Aspen has a perfect software for the calculation of all the utilities and the integration of all the heat

exchange units in the process, Aspen Energy Analyzer.

Aspen Energy Analyzer allows the user to see a detailed heat exchanger network (HEN)

diagram and the composite curves used in generating the heat integration. The HEN diagram shows

heat exchanger pairings and approach temperatures for the streams. The Composite curves can show

us the Pinch lines [52].

3.4 Activated Exchanger Design & Rating

To specify the Activated Economics, the software of Activated Exchanger Design & Rating

(EDR) has been used.

With this activated analysis tool available in the AspenTech’s V8.4 simulators Aspen Plus® and

Aspen HYSYS®, helps process engineers quickly evaluate feasible design alternatives and

immediately assess the impact on flowsheet capital and operating expenditures. The software helps to

convert simple heat exchanger models to rigorous models, and provides quick access to rigorous

model details. The software makes perfectly us of the Shell-and-Tube Heat Exchangers (STHE)

classification, TEMA shown in figure 17. The Tubular Exchanger Manufacturers Association (TEMA) is

trade association of leading manufacturers of shell and tube heat exchangers, who have pioneered

the research and development of heat exchangers for over sixty years [53]. Beside the TEMA

designations, the tube layout, baffling, pressure drop and mean temperature difference are to be

mentioned. The EDR dashboard uses visual indicators to convey information about heat exchangers’

model status and shows all operational risks, if any, that have been detected in rigorous heat

exchanger models [54].

28

Figure 17: TEMA Designation System Example [55].

29

4 Modeling

4.1 Process simulation

Off all the possibilities to reuse carbon oxide, the conversion to methanol was chosen. The

optimization of this conversion is based on the methanol simulation out of the literature [50] . This

simulation is based on the recovery of CO2 produced in the Cimpor cement factory in Lisbon. As the

cement production process in all cement factories around the world are responsible for an insignificant

amount of carbon dioxide discharge, the capture and conversion to a useful product as methanol is of

great interest. The alternatives for carbon dioxide reuse are still in a pilot stage or need more

investigation. Geologically wise, Carbon storage is not possible in the area around Lisbon or at the

Cimpor factory. With this in mind, the conversion of carbon dioxide into methanol is an ideal solution

for the emissions of Cimpor. Methanol has many applications, a wide market and is not hard to

transpose.

Once the flue gasses of the cement factory have been captured, the conversion to methanol

only needs an excess of hydrogen to ameliorate the reaction with CO2. The following downstream

process consists of separating mainly the not reacted hydrogen, oxygen, nitrogen and alternative

products as methane, dimethyl ether and water.

The process that is simulated is based on two inlet streams, a pure hydrogen stream and a flue

gas stream. The flow and composition of the flue gas stream were copied from the literature

[50],shown in table 1, which are based on a document of the European Commission on the best

available techniques in cement, lime and magnesium manufacturing industries [56].

Table 1: Composition of the flue gas feed.

Components

of Cement flue gas

Amount of each components

per tonne of cement produced

Mass (kg) Mole

CO2 672 5.4

O2 389 4.3

N2 1975 24.9

H20 132 2.6

TOTAL 3168 37.2

30

The goal of the different scenarios, simulations in Aspen plus®, is to optimize several

parameters of the CO2 to methanol production process. These parameters will indicate what

optimization is needed. The parameters are based on the productivity and energy consumption of the

process, and will be the key performance indicators (KPI) of this research, shown in table 2. In

combination with the results of the Aspen plus® tools, Activated Economics, Aspen Energy Analyzer

and Activated Exchanger Design & Rating, it is possible to consider if the carbon dioxide reuse

technology is valuable.

Table 2: KPI's descriptions.

KPI Description

Hydrogen Consumption

Displays the amount of hydrogen that is fed to the

reactor to produce methanol, expressed in kg H2 / kg

CH3OH

(𝐻2)𝑐𝑜𝑛𝑠. =(𝐻2)𝑓𝑒𝑒𝑑

(𝐶𝐻3𝑂𝐻)𝑝𝑟𝑜𝑑𝑢𝑐𝑒𝑑 (13)

Hydrogen efficiency

Shows how many hydrogen of the feed has reacted

with the carbon dioxide, expressed in %

(𝐻2)𝑒𝑓𝑓. =(𝐻2)𝑓𝑒𝑒𝑑 − (𝐻2)𝑝𝑢𝑟𝑔𝑒

(𝐻2)𝑓𝑒𝑒𝑑 (14)

Productivity

Gives the mass of methanol that is produced per

kilogram carbon oxide feed. Given in kg CH3OH /

CO2

𝑃𝑟𝑜𝑑𝑢𝑐𝑡𝑖𝑣𝑖𝑡𝑦 =(𝐶𝐻3𝑂𝐻)𝑝𝑟𝑜𝑑𝑢𝑐𝑒𝑑

(CO2)𝑓𝑒𝑒𝑑 (15)

Electrical power consumption

Presents the consumption of electric energy by the

equipment for a methanol production of one hour, in

kWel/hr.

𝐸𝑙𝑒𝑐𝑡𝑟𝑖𝑐 𝑝𝑜𝑤𝑒𝑟 =∑𝐸𝑙𝑒𝑐𝑡𝑟𝑖𝑐 𝑑𝑢𝑡𝑖𝑒𝑠


Thermal consumption

Presents the consumption of thermal energy by the


kWth/hr.

𝐸𝑙𝑒𝑐𝑡𝑟𝑖𝑐 𝑝𝑜𝑤𝑒𝑟 =∑𝐻𝑒𝑎𝑡 𝑑𝑢𝑡𝑖𝑒𝑠


Cooling water consumption

Presents the consumption of cooling water by the


kg/hr.

𝐸𝑙𝑒𝑐𝑡𝑟𝑖𝑐 𝑝𝑜𝑤𝑒𝑟 =∑𝐶𝑜𝑜𝑙𝑖𝑛𝑔 𝑤𝑎𝑡𝑒𝑟

(𝐶𝐻3𝑂𝐻)𝑝𝑟𝑜𝑑𝑢𝑐𝑒𝑑

(

18)

CO2-emmisions

Expresses the amount of carbon dioxide that has not

reacted.

(CO2)𝑓𝑒𝑒𝑑 − (CO2)𝑝𝑢𝑟𝑔𝑒(CO2)𝑓𝑒𝑒𝑑

(19)

31

4.2 Methanol Production Simulation

4.2.1 The baseline scenario

The conversion of carbon dioxide to methanol is just a hydrogenation of carbon dioxide that is

widely described in many articles. In general the conversion consists of three main reactions,

(equation 6, equation 9 and equation 10) and two side reactions (equation11 and equation12),

depending on the selectivity of the catalyst [57]. Besides the selectivity, the catalyst is responsible for

the conversion in the reactor. Among the choice of catalyst, the conditions in the reactor where this

catalyst has to work in are very important and decisive for the process. Generally heavy conditions,

high pressure and temperature [40], are necessary to have a high selectivity and conversions to

methanol. With these conditions it is favourable to have an excess of hydrogen as shown in figure 18.

Figure 18 shows different selectivity’s and conversions under equilibrium at different conditions of feed

CO2:H2 ratio and pressure as a function of temperature.

Figure 18: Methanol selectivity (SMeOH) at equilibrium (left) and Equilibrium conversion of CO2 (XCO2) (right) at different conditions of feed CO2:H2 ratio and pressure as a function of temperature [40].

The baseline scenario, shown in figure 16 (appendix 1), will be working under the same

conditions as the simulation in the literature [50]. The study will not be on the conversion or selectivity

of the reaction, CO2 to methanol, but on the optimization of the KPI and the heat integration of this

simulation. Therefore the conditions, conversion and selectivity’s are copied, coming from the

following reference [40].

32

Table 3 presents the selectivity and conversion in the reactor, under a pressure of 330 bar, a

temperature of 260 °C and an excess of hydrogen feed of CO2:H2 of 1:10.

Table 3: Carbon dioxide conversion and selectivity's at 260°C and 330 bar

CO2 conversion (%)

Selectivity (%)

CO CH4 CH3OH CH3OCH3

95.3 1.4 O0.3 98.2 0.1

The reactor that is employed, stoichiometric reactor (figure 19) is known as a type of reactors

where reaction kinetics are unknown or unimportant, but stoichiometry and extent of reaction are

known. In addition, stoichiometric reactors can perform product selectivity and reaction heat

calculations.

Figure 19: Stoichiometric reactor in Aspen Plus®.

The occurring reactions are filled in the simulation together with its fractional conversions, for

example methanol shown in figure 20. It is worthy to notice that equation 10 hasn’t been mentioned in

the reactor, so no methanol is synthesized from produced carbon monoxide.

33

Figure 20: Methanol reaction implementation in the stoichiometric reactor interface.

The feed streams, pure hydrogen (108kg/hr) and the flue gas (1120 kg/hr) from the cement

factory are compressed to 330 bar and afterwards cooled to 260°C.

The downstream process counts four separating steps, using two flash units and 2 distillation

towers. The product stream first goes through a Flash2 unit that separates most of the gas phase,

especially hydrogen, from the liquid phase of the product stream. The flash unit operates under 50 bar

and 25°C, involving a 98% recovery of methanol. The flash is followed by a distillation column, which

serves to separate the remaining gasses, resulting in a 95% mass recovery of methanol. The gaseous

distillate stream continues in the second Flash2 unit to recover methanol by cooling it down till 0 °C.

The liquid phase methanol from the bottom stream of the second flash unit gets mixed with the bottom

stream of the first distillation column. The mixed stream consists of more or less 60% water. The final

column therefor is designed to separate the water from the methanol. With a mass recovery of 95% of

methanol, the distillate stream has 97.3 mole fraction of methanol.

The stream results of the simulation are used to calculate the KPIs. The goal of this research is

to improve these indicators by modifying the simulation. The KPIs of the first simulation is shown in

table 4, these KPIs will be the base line of this work.

34

Table 4: KPI for the baseline simulation

Key Performance Indicators

Hydrogen consumption 0.69 kg H2 / kg CH3OH

Hydrogen efficiency 28.6%

Productivity 0.659 kg CH3OH / kg CO2

Electrical power consumption 1257 kWel

Thermal power consumption -1407 kWth

Cooling water consumption 1730 kg/hr

CO2- emissions 17.6 kg/hr

The main improvements will be the reduction of hydrogen consumption, lowering the electrical

and thermal power and trying to use as less possible cooling water. The cooling water will be

converted into high pressure steam and can be integrated in the process or be sold.

4.3 The integration process

4.3.1 Scenario 2

The CO2 to methanol process is improved with Aspen Plus® throughout this thesis in 9

scenarios, starting from the baseline simulations of previous work [50]. The problem was that the

simulation was not realistic enough in terms of heat exchangers and compressors. The baseline

simulation compressed the feed streams in one compressor from 1 to 330 bar. Those conditions are

only feasible with extraordinary compressors, which are very expensive and consume a lot of energy.

One of the main goals was to lower down the electrical power consumption. To obtain a more

economical friendly process, the compression process will undertake multiple stages. The

compression process is shown in figure 21; each compressor unit is followed by a heat exchanger

[58].

Figure 21: Flowsheet compression unit [58].

35

In scenario two (appendix 2) the multiple compressors were introduced. Each compression was

of a factor 4 followed by a heat exchanger, except the fifth compression unit was from 256 bar to 330

bar. This multiple stage compression results in total electrical power consumption of 800.5 kilowatt

(kWel), a reduction of almost 36%. Actually all the consumption KPI’s, except hydrogen, were reduced.

Table 5: Reduced KPI in simulation 2.

Key performance indicator

Electrical power consumption 800.5 kWel

Thermal power consumption -950.5 kWth

Cooling water consumtion 1109 kg/hr

4.3.2 Scenario 3

In Scenario three the focus is on the consumption of cooling water and the production of super-

heated steam (appendix 3). The main change in the simulation is the connection of all the cooling

streams. Every outlet cooling stream will be the inlet cooling stream of the next heat exchanger. Figure

22 shows the connection of cooling water streams in the hydrogen compressing section (CW-H2). To

achieve superheated steam, the inlet cooling water was compressed to 5 bar at 20 °C. Making use of

design specifications, optimizations and constraints, the flows of the cooling streams were decreased

and totally vaporized after the fifth heat exchanger.

The overall cooling water consumption decreased to 1035 kg/hr, a reduction around 7%. The

most valuable part of this simulation is the superheated steam that is achieved at the end of both

compressing units, it will be important for the heat integration in the total CCUS process.

Figure 22: Cooling water connection for hydrogen feed, simulation 3.

36

4.3.3 Scenario 4

One KPI that stands out is the efficiency of hydrogen, only 28.6% of the hydrogen reacts with

carbon dioxide. Having in mind that the production of hydrogen is very expensive through the electrical

power consumption of the electrolysis, a recycle of the hydrogen will be necessary. The hydrogen is

separated from the main product stream at the first flash unit. Around 70% of the feed has not reacted

with carbon dioxide and must be reused. The problem is that the stream GAS-S1 from scenario four

(appendix 4) has a lot of impurities and a different pressure than the inlet stream H2-IN. A valve was

installed to bring back the pressure to one bar, followed by a palladium membrane separator that can

achieve a hydrogen purity of 99.9999 % [59]. In the scenario 4, around 99% off the hydrogen got

recovered, which under pilot stage conditions could not be realized, further research has to be done

on the recovery rate. After the separation the 99.9999% pure hydrogen got mixed with the pure feed

stream, shown in figure 23.

Figure 23: The hydrogen recycle simulation 4.

The results are not that relevant, due the almost perfect recovery of hydrogen. The 99%

recovery means 99% hydrogen efficiency or a hydrogen consumption of 0.20kg hydrogen for each kg

of methanol. In literature was found that the performances of a palladium membrane in an industrial

configuration: considering a feed made of 50% H2 and 50% CO2 by volume (V%), a total flow-rate of

0.016 mole a second (mol/s) and with a membrane length of 3 m, the hydrogen recovery is higher

than 90% with a product purity over 99.88% [60]. That would refer to consumption around 0.25 kg of

hydrogen for the production of one kg methanol.

4.3.4 Scenario 5 and 6

Scenario five and six (appendix 5 and appendix 6) relates to the integration of the produced

steam in the compressing units. The super-heated steam will be integrated in the reboilers of the two

distillation towers, simulation five on the first distillation column and simulation 6 on the second

distillation column. Due a sensitivity test, made in two other simulations (appendix 9), the exact

amount of super-heated steam for the reboiler was calculated. Employing the exact amount of super-

heated steam, splitters, heat exchangers and the reflux ratios of the distillation columns, the reboil

37

process of the distillation towers was copied without external steam or a reboiler. Figure 24 shows the

integration of the superheated steam of the hydrogen compression section (CW-H2-out) to warm up

the reflux stream (S2) of the distillation column (DEST-1) in heat exchanger (WW-in). The excess of

steam can be employed in the carbon capture process or can be converted in a electricity by a turbine

(TURB1 and TURB2), as is the case in this simulation shown in figure 24.

Figure 24: Super-heated steam integration simulation 5 and 6.

This adaption mainly affects the thermal power consumption as the reboilers of the two

distillation columns are the only steam consumers in the process. The thermal consumption for steam

drops to zero. The overall thermal consumption decreases to -846.4 kW th. The excess of steam that is

converted into electricity is around 40.75 kWel. The electricity can be reused in the compressors so

that the overall electrical power consumption drops to 775.5 kWel.

4.3.5 Scenario 7 and 8

In scenario seven (appendix 7) the recycle loop of simulation 4 was adapted. Instead of

dropping the pressure of GAS-S1 back to 1 bar, it was recycled on 50 bar. Therefore the simulation

needed some modifications. Firstly valve 3 after the Separator 1 was removed so the pressure stayed

on 50 bar, followed by adjusting compressor H2-CMP-3 to 50 bar instead of 64 bar.Therefore the

cooled feed stream H2-50-XO could be mixed with pure recycle stream RECYCLH2 on 50 bar. The

direct recycle simulations of 50 bar had some warnings, to solve the warnings a new stream was

introduced RECYH2IN with the exact same specifications as the RECYCLH2 and mixed with H2-50-

XO. The stream that will enter compressor 4, H2-BC-4, has the exact same mass flow as if there

would be a direct recycle loop, but the simulation works without any warnings as pictured in figure 25.

38

Figure 25: H2 recycle on 50 bar simulation.

By recycling the excess of hydrogen on 50 bar, means that the overall electrical energy

consumption due the compression unit will drop. Specifically the first three compression units only

need to compress 31.5 kg/hr instead of 108 kg/hr. This decrease in hydrogen flow also affects the

other consumption indicators, except the hydrogen consumption, shown in table 6. With regard to the

compressors the reduction in electrical power consumption is 202 kWel compared with scenario 6

which is around 26% and almost 684 kWel for the baseline simulation that is approximately 55%.

Table 6: reduced KPI in simulation 7.

Key Performance Indicator



Cooling water consumption 736.6 kg/hr

Scenario eight (appendix 8), is the final simulation and concludes all the modifications that were

made in the previous simulations stating from the base line simulation. Figure 26, shows the final

result of all the adjustments together.

39

Figure 26: Flowsheet Scenario 8.

40

Because of the decrease in hydrogen consumption the total amount of cooling water has as well

decreased. Less cooling water means less super-heated steam. What makes us obliged to combine

the two super-heated streams to have enough steam for both distillation columns, shown in figure26.

The excess of steam can be converted to electricity in a turbine and reused (scenario 8), employed in

the CCS-process or be sold. The overall changes in the optimization process are summarized in table

7. It is obvious that mainly the consumption indictors have changed.

Table 7: KPI after the all the modifications, simulation 8.

Key Performance Indicators

Hydrogen consumption 0.20kg H2 / kg CH3OH

Hydrogen efficiency 97.4%

Productivity 0.665 kg CH3OH / kg CO2



Cooling water consumption 758.7 kg/hr

CO2- emissions 17.6 kg/hr

The cooled high pressure steam that was used in the heat exchanger of the second distillation

(WW-D2) tower, to replace the reboiler, is now in liquid phase and can be used to cool down the

stoichiometric reactor (scenario 9). The reactor has to be cooled to keep the reactor conditions steady,

260°C. The cooling is necessary because the carbon to methanol reaction is exothermic.

Unfortunately this was not possible to simulate, but the calculation were done. The heat duty of the

reaction is 83 kWth. The reactor will need 137.1kg/hr of cooling water in the condition that stream CW-

GO leaves WW-D2. This is possible because flow of CW-GO is 179 kg/hr. Considering this

integration, the thermal power consumption will again reduce with 83kW th/hr. This brings the total

thermal power consumption to 520.6 kWth and total cooling water to 646.5 kg/hr.

4.4 Economic simulation

To have an idea of the costs of the simulation in real life, Aspen has the Activated Economics

tool. This is used to measure all the equipment cost, utilities and weight, shown in figure 27. Mainly the

costs of the compressors, distillation towers and membrane separator were done with this tool.

41

Figure 27: Activated Economics simulation 8.

The calculations of the heat exchangers were done by the Activated Exchanger Design &

Rating software. This software makes it possible to create and size the heat exchangers for whatever

professional qualifications that are needed. The specific properties of each heat exchanger are

summarized in its TEMA sheet (appendix 11).

In scenario 8 there are twelve heat exchangers, ten for the compressing unit and two to replace

the reboilers in the distillation towers. All have been sized by Activated Exchanger Design & Rating, as

for example on figure 28. The heat exchangers until 50 bar used the B.E.U.(appendix 10). TEMA

design: Bonnet (integral cover), one-pass shell and a u-Tube bundle. This design is the cheapest

possible in low and medium pressure conditions. All the heat exchangers that are working in

conditions higher than 50 bar (high pressure conditions) are using the D.E.U (appendix 11). TEMA

design code, Special High-pressure Closures, one-pass shell and a u-Tube bundle.

42

Figure 28: Activated Exchanger Design &Rating of WW-gas-1.

The Activated Exchanger Design & Rating software also calculates the cost and weight of the

heat exchanger, as illustrated on figure 29, which are useful to calculate the capital costs of the total

process.

Figure 29: Costs and weights of WW-gas-1.

On the TEMA sheets of the heat exchangers there is much more information to be fiend, but

this information relies more to mechanical engineers than chemical engineers. All the TEMA sheets

are to be fiend in the appendices.

43

Aspen Energy Analyzer software is used to measure the several utilities and in how far the

integration for the heat exchangers has succeeded. Once utilities are known, the operating cost can

be calculated.

Figure 30 presents all the processes of heat exchange that are going on in the final scenario 9.

It summarizes the duties, temperatures of the streams, the recoverable duties and more. By looking to

the recoverable duty, it seems that the optimization was successful. The only processes where heat

can be recovered are the WW-GAS-1, WW-H2-1 and stoichiometric reactor, the heat duty that occurs

at the SEP1 and SEP2 (flash2 units) are the result of the phase changes that are necessary to fulfill

the separation process. The duties can’t be exchanged with other streams. For the stoichiometric

reactor there has already been found a solution, the use of stream CW-GO. This liquid phase cooling

stream gets vaporized by the exothermic reaction heat in the reactor and can be reused or be sold as

steam. Of the five recoverable duties, only WW-GAS-1 and WW-H2-1 can be optimized. Still the

recoverable duty is only 0.008832Gcal/hr or 10.3kWth, not really worthy to recover. The condenser of

the second distillation tower has no values, because Aspen doesn’t consider the condenser as a heat

exchanger. Still the duty is known out of the thermal results of the distillation column, 118.2 kWth.

44

Figure 30: Heat exchanger details by Aspen Energy Analyzer.

45

The Aspen Energy analyzer software can also visualize all the stream data in a grid diagram,

figure 31. This projection of the heat exchange of the simulation is used in general to start the pinch

analyzing. The condenser is pictured as the red dotted line. Hot streams in red and cold streams blue.

All the process heat exchanger are shown with grey dots, all the coolers for as result of phase

changes (dummy coolers) are shown with blue dots and all the heaters for phase changes (dummy

heaters) are shown with red dots. The dummy coolers and heaters have no real purpose, because you

can’t transfer the heat and that’s why the heat duties are equal opposites of each other. Except for the

case of the stoichiometric reactor, there the HOTWATER utility is used to calculate the heat transfer.

Figure 31: Grid diagram in Aspen Energy Analyzer.

The problem is that the optimizations that the process has been through resulted in threshold

problems. The characteristics of these problems are that these processes either need hot utility or cold

utility and not both. Threshold problems can be divided into two broad categories for purpose of

design. In the first type, the closest temperature approach between the hot and cold composites is at

the “non-utility” end and the curves diverge away from this point. The second type, there is an

intermediate near-pinch (pseudo pinch), which can be identified from the composite curves, figure 32,

as the region of close temperature approach. Now the question is whether pinch design method

(PDM) should be modified to deal with design of the threshold problems [61]? From the composite

curves, figure 32, it is also very clear that the process requires no hot utility and the pinch point is at

the hot end. However, the simulation exhibits a pseudo pinch as shown in figure 32.

46

Figure 32: Hot and Cold composite curves of simulation 8.

For the present process hot utility requirement is 0kW whereas the cold utility requirement is

514 kWth. The pseudo pinch and pinch are illustrated on the grand composite curve, shown in figure

33. The pinch is situated around 400 °C and the pseudo pinch is about 160 °C.

Figure 33: Grand composite curve for Simulation 8.

Thus, in this case, it is to treat the threshold problem as a pinch problem and pseudo pinch as

pinch and design away from the pseudo pinch. The only difficulty in treating this problem as pinch

problem is that one half of the problem (the hot end as shown in figure 32) will not offer any flexibility

to match against the hot utility [61].

47

The likely reason that there is a threshold problem in this simulation is because of substitution of

the reboilers in the distillation towers for heat exchangers.

The composite curves also perfectly illustrate the areas where there are phase changes. All the

horizontal parts of the blue line refer to the cooling water in the compressions unit that gets vaporized.

Al the horizontal parts of the red line refer to the vapor streams that changes to liquid for example in

the WW-D1 and WW-D2 as the flash units. [61].

49

5 Results

To summarize all the results of the simulations, KPI and economic analysis, some useful

graphics are made. The last change was the reuse of the cool water in stoichiometric reactor. The KPI

improvement between the baseline and scenario 9 is summarized in table 8.

Table 8: The KPI improvement between the baseline and scenario 9.

Key Performance Indicator

Units Base line

scenario Scenario 9

Improvement

(%)

Hydrogen consumption kg H2 / kg CH3OH 0.69 0.20 71

Hydrogen efficiency % 28.6 97.4 68.8

Productivity kg CH3OH / kg CO2 0.659 0.665 0

Electrical power consumption kWel 1257 553.6 56

Thermal power consumption kWth -1407 -520.6 63

Cooling water consumption kg/hr 1730 646.5 62

CO2- emissions kg/hr 17.6 17.6 0

The KPI optimizations are well illustrated in the following graphics. Every graphic represents a

certain KPI for the eight or nine different scenarios. In addition of the appendices, the whole evolution

of the carbon reuse process is visualized.

Figure 34: Hydrogen consumption over the 8 simulations.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0 1 2 3 4 5 6 7 8

H2 c

on

sum

pti

on

(k

g H

2/kg

MeO

H)

Scenario Number

simulation 1 to 8

50

Figure 34 shows that the optimization of the hydrogen consumption is very good because of the

almost perfect hydrogen recuperation process. The membrane was set on 99% recovery of the

hydrogen, which is kind of unrealistic. This results in almost no loss of hydrogen in the process what is

translated in very high hydrogen efficiency, shown on figure 35.

Figure 35: Hydrogen efficiency.

Hydrogen is very expensive, so this recycling of hydrogen will be a must. If the goal is to

produce methanol in green way, the production of hydrogen can’t be done by combustion natural

resources. The necessary electricity, 4.8 kWhel/Nm3, for the electrolysis of water will have to come

from environmental friendly processes [58]. It can be produced by wind energy, biofuels, sun energy

or many other ways. Problem for Cimpor is that there is no directly available area around the site to

install these environmental friendly processes. Luckily hydrogen is easy to transport so it can be

produced somewhere else. If the electrolysis is on the plant, the production cost of the hydrogen can

be decreased through selling the pure oxygen generated by water electrolysis [58].

Figure 36 shows the decrease of electrical power consumption.

Figure 36: Electrical power consumption over the 8 simulations.

0

20

40

60

80

100

0 1 2 3 4 5 6 7 8

H2 E

ffic

ien

cy

(%)

Scenario Number

simulation 1 to 8

0

200

400

600

800

1000

1200

1400

0 1 2 3 4 5 6 7 8

Ele

ctri

cal P

ow

er

Co

nsu

mp

tio

n (k

We

l/h

r)

Simulation Number

simulation 1 to 8

51

It is clear that the use of different compressions units is responsible for a drop of more than 400

KW/hrel. The other drop in consumption is relied to the recycle of hydrogen on 50bar. Finally the total

electrical power consumption got to 553.6 kWel or 3.5kWel per kilogram of methanol.

The thermal power consumption got measured in two different ways, one by the just counting all

the heat duties of every equipment and one way by the Aspen Energy Analyzer. In the end there was

a difference of around 100 kWth, this is because Aspen Energy Analyzer doesn’t recognize the

condenser as a equipment that exchanges heat.

Figure 37: Thermal power consumption simulation 1 to 8.

Figure 37 shows all the heat duties of all heat exchange equipment’s. Because of the heat

integration modification, the process can be illustrated as a scenario without heat production in the

reboilers or without cooling in the reactor. Both ways of calculation give thermal power consumption

around 725 kWth before the heat integration, scenario 7. In scenario 9 the thermal power consumption

dropped to 520 kWth, which is around 63%.

Figure 38 represent the consumption of cooling water. In this scenario the use of multiple

compressors is better for the cooling water consumption. The recycle of the hydrogen at 50 bar gives

a reduction of cooling water as well. The third drop in the graphics is due the reuse of the cooling

water in the stoichiometric reactor. The final consumption of cooling water is 646.5 kg/hr, a reduction

of 62%.

-675

-600

-525

-450

-375

-300

-225

-150

-75

0

75

150

1 2 3 4 5 6 7 8 9

Th

erm

al P

ow

er

Co

nsu

mp

tio

n (k

Wth

/hr)

Scenario Number

HX-GAS

HX-H2

R-stoic

reboiler d1

reboiler d2

condensor d2

52

Figure 38: Cooling water consumption simulation 1 to 9.

The two last KPI, CO2-emissions and productivity didn’t change. The scenario represents an

effective solution to convert carbon dioxide into a more valuable product, methanol. The process is

also perfectly possible for the conversion of carbon dioxide into methane or dimethyl ether, depending

on the demand of the market [50].

Heat exchanger network design was based on Pinch analysis performed with Aspen Energy

Analyzer. The Activated Exchanger Design & Rating software of aspen sized all the used heat

exchangers; the result is shown in figure 39. It is obvious that loads of money can be saved if the

working conditions would be less hard. Imagine working under 150 bar, that would save more than

100 000 dollars in production costs of heat exchangers, what to say about the costs of the

compressors. These conditions exist in literature, carbon dioxide to methanol scenario’s working at a

pressure of 78 bar and a temperature of 80°C [58].

0

200

400

600

800

1000

1200

1400

1600

1800

0 1 2 3 4 5 6 7 8 9

Co

olin

g W

ate

r C

on

sum

pti

on

(k

g/h

r)

Scenario Number

simulation 1 to 9

53

Figure 39: Heat exchangers size.

To have an idea of the total capital cost of the carbon capture and reuse plant, the cost of the

carbon capture in the cement industries has to be included. From previous research the costs of the

two possible carbon capture technologies, Oxy-combustion and post-combustion capture are

implemented in this study. Oxy-combustion has the lowest cost solution for CO2 capture at new-build

cement plants but research and development is needed to address a number of technical issues to

enable this technique to be deployed. Costs are estimated to be €34.3 per tonne of CO2 captured for a

1 Mt/y European cement plant [62]. The estimated costs of post-combustion capture are substantially

higher: €59.6/ per tonne of CO2 captured 1 Mt/y European cement [62]. The cost of CO2 capture at a

cement plant using oxy-combustion is expected to be similar to the cost at a typical coal-fired power

plant. The quantity of oxygen required per tonne of CO2 captured is about three times lower at a

cement plant but the economies of scale are less favorable. The cost of post-combustion capture at a

cement plant is expected to be substantially higher than at a power plant [62]. Table 9 summarizes

the cost off the carbon capture in the cement industries.

0

20

40

60

80

100

120

0

100

200

300

400

500

600

700

800

900P

rod

uctio

n C

ost

(USD

x103)

He

at E

xch

ange

r W

eigh

t (k

g)

Weight

Production Cost

54

Table 9: Summary of Cement Plant Costs With and Without CO2 Capture [62].

Unit Base case Post combustion

Oxy-combustion

No capture Capture

Capital cost €M 236 558 327

Operating costs

Fuel €M/y 6.7 21.5 6.9

Power €M/y 4.0 -1.1 8.7

Other variable operating costs €M/y 6.1 10.6 6.4

Fixed operating costs €M/y 19.1 35.3 22.8

Capital charges €M/y 29.7 63.1 36.9

Total costs €M/y 65.6 129.4 81.6

Cement production cost €/t 65.6 129.4 81.6

CO2 abatement costs €/t - 63.8 16.0

Cost per tonne of cement product

Cost per tonne of CO2 captured €/t - 59.6 34.3

Cost per tonne of CO2 emissions avoided €/t - 107.4 40.2

In addition there are the costs of the carbon reuse process, that is simulated on Aspen plus

®

and calculated in Aspen Activated Economics analyzer for the capital costs and for Aspen energy

analyzer for the utilities. The results of the carbon reuse simulation of scenario nine are summarized

and compared with base line scenario in table 10.

Table 10: Economical evaluation of methanol plant scenario 9.

Unit Base line scenario Scenario 9

Capital cost €M 12.9

Operating costs €M/y 4.42 1.49

Total raw materials €M/y 3.1 0.908

Total utilities cost €M/y 1.32 0.582

Sales of methanol €M/y 505 0.505

Margin €M/y -5.2 -2.27

Cost per tonne methamol product

Capital cost per tonne €/t 929 929

Operating cost per tonne €/t 3191 1076

Total raw materials €/t 2236 656

Total utilities cost €/t 955 421

Sales of methanol €/t 365 365

Margin €/t -1641

55

The only raw material is hydrogen, as the carbon dioxide is a flue gas from the cement plant.

Hydrogen has a distribution cost of 3.7$/kg or 3.27 €/kg. The National Renewable Energy Laboratory

(NREL) analyzed the cost of hydrogen production via wind-based water electrolysis at 42 potential

sites in 11 states across the United State of America. This analysis included centralized plants

producing the Department of Energy (DOE) target of 50,000 kg of hydrogen per day, using both wind

and grid electricity [63]. The utilities are cooling water at 0.065 €/t (highly dependent on the type of

industry) and electricity has an average cost in Portugal that is around 0.12 €/kWh [64]. The selling

price of methanol is around 0.365$/kg or 3.27€/kg [65]. The calculations have been done as

explained in 3.2 The Activated Economics. The margin is calculated with equation 20.

𝑀𝑎𝑟𝑔𝑖𝑛 = 𝑠𝑎𝑙𝑒 𝑜𝑓 𝑚𝑒𝑡ℎ𝑎𝑛𝑜𝑙 − (∑ 𝑐𝑎𝑝𝑖𝑡𝑎𝑙 𝑐𝑜𝑠𝑡

10+∑ 𝑜𝑝𝑒𝑟𝑎𝑡𝑖𝑛𝑔 𝑐𝑜𝑠𝑡) (20)

Out of the economics evaluation it is clear that the costs have been improved over the

scenarios, the improvement is shown in table 11.

Table 11: Economic improvement.

Economic

Improvement

Operating costs 66.3%

Total raw

materials 70.1 %

Total utilities cost 56%

Sales of methanol 0%

Out of the economic analyze for the whole plant (cement, carbon capture and carbon reuse),

there can be concluded that the carbon capture process is only a fraction of the cost of the carbon

reuse process. A reason can be the small scale simulation (feed flows are really low) compared to the

1 Mt/y European cement plant [62], resulting in a higher capital cost per kilogram methanol produced.

But the main reason is the higher operating costs, due the hard reaction conditions 330 bar and 260

°C and practically due the consumption of hydrogen. The hydrogen counts for 61% of the operating

cost and would probably be higher if there wasn’t a 99% recovery of hydrogen, due the membrane

separator. In total the balance of the methanol carbon reuse is negative and will probably continue to

be if there are no more efficient manors to convert carbon into methanol or producing hydrogen in a

cheaper way. The only productive way is to produce hydrogen with renewable energy, economically

and environmentally.

Figure 40: Cost tonne methanol ($/t).

56

6 Conclusion and Future work

To argue the feasibility of CO2 conversion processes, nine different improvement scenarios on

the design of carbon dioxide conversion to methanol process were made. These improvements were

partially duty optimization and partially heat integration. These improvements resulted in a decrease of

KPI between 50-70%. The Aspen plus® software was a perfect tool for the optimization of the utilities

and the heat integration of the 12 heat exchangers in the process. The result of this integration was

showed in the Aspen Energy Analyzer, only 2.8kWth is recoverable. The grid diagram in Aspen Energy

Analyzer concluded that the process has a threshold problem. The characteristic of this problem was

that there is only cold utility needed, 514 kW th. The threshold problem is assumed to be caused by the

substitution of the reboilers in the distillation towers by heat exchangers, whom are feed by the

superheated steam (cool water) coming from the compressing unit. Out of the grand composite curves

the pinch is situated around 400 °C and the pseudo pinch is about 160 °C. The heat exchangers were

sized as economically possible with the help of The Activated Exchanger Design & Rating software.

The low pressure and medium pressure heat exchangers have B.E.U. TEMA design code and the

high pressure heat exchanger have D.E.U. TEMA design code. The total cost of the heat 12 heat

exchanger was 190 thousand euro, future work will be necessary to check these prices. Together with

the rest of the equipment the total Capital cost was around 13 million euro. The Operating cost was

roughly 1.5 million euro and especially due the hydrogen consumption. In the final scenario nine, the

cost of one tonne of methanol is 1.6 thousand euro for a 238 kg/hr CO2 feed stream. Assuming

working with higher feed flows, the impact of the capital cost will decrease and the main cost will be

hydrogen [50]: Large demand of hydrogen is expected with higher flows, which is both expensive and

laden with sizable CO2 footprint, is the biggest hurdle to achieve a more sustainable solution for the

methanol production via CO2 hydrogenation. There are various renewable energy sources to

considerate): biomass, hydroelectric power, nuclear power, photovoltaic, solar thermal power, wind

power, etc. Hydroelectric power and biomass based hydrogen production processes are nowadays

the most interesting with a methanol production cost of 350 euro a tonne. Most likely, the wind and

solar thermal power based H2 production options can also become economically feasible if their

power generation costs decrease further in the future [66]. Good to know is that methanol plants are

not limited to places where renewable energy is available for CCS. Hydrogen can be generated in

places where renewable energy is available and then exported to methanol synthesis plants [58]. The

hydrogen membrane separator´s conditions will have to be defined in the future, to make the research

more realistic.

Although this CO2 to methanol process needs further research and development, the future of

this technology looks bright. Global methanol demand will rise significantly – from 60.7 MMt in 2013

to109 MMt by 2023. More than 50 MMt of new methanol capacity has been announced [67].

57

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63

Appendices

A1 - Scenario 1

Figure 41: Flowsheet scenario 1.

64

A2 - Scenario 2


65

A3 - Scenario 3


66

A4 - Scenario 4


67

A5 - Scenario 5


68

A6 - Scenario 6


69

A7 - Scenario 7


70

A8 - Scenario 8


71

A9 - Scenario 8 – Stream Tables

BOT-D1 BOT-D2 CW-5BAR CW-G-1 CW-G-2 CW-G-3 CW-G-4

SPLIT2 SPLIT4 TURB1 WW-GAS-2 WW-GAS-3 WW-GAS-4 WW-GAS-5

DEST-1 DEST-2 MIX-CW-O WW-GAS-1 WW-GAS-2 WW-GAS-3 WW-GAS-4

LIQUID LIQUID VAPOR LIQUID MIXED MIXED VAPOR

Substream: MIXED

Mass Flow kg/hr

CO2 2.82E-11 9.64E-09 0 0 0 0 0

O2 2.82E-12 9.39E-14 0 0 0 0 0

N2 5.03E-12 1.50E-13 0 0 0 0 0

H2O 1.615.786 3.317.601 2.822.842 3.250.000 3.250.000 3.250.000 3.250.000

CO 2.07E-14 6.45E-16 0 0 0 0 0

METHANOL 1.766.720 0.3840464 0 0 0 0 0

H2 2.54E-19 9.48E-22 0 0 0 0 0

METHANE 3.08E-13 1.91E-14 0 0 0 0 0

DME 1.29E-04 3.00E-04 0 0 0 0 0

Liq Vol 60F cum/hr

CO2 3.43E-14 1.17E-11 0 0 0 0 0

O2 4.72E-15 1.57E-16 0 0 0 0 0

N2 9.61E-15 2.88E-16 0 0 0 0 0

H2O 0.16189 0.3323995 0.2828283 0.3256264 0.3256264 0.3256264 0.3256264

CO 3.96E-17 1.23E-18 0 0 0 0 0

METHANOL 0.2223943 4.83E+01 0 0 0 0 0

H2 6.76E-21 2.52E-23 0 0 0 0 0

METHANE 1.03E-15 6.39E-17 0 0 0 0 0

DME 1.97E-08 4.57E-08 0 0 0 0 0

Total Flow kmol/hr 1.448.271 1.842.747 1.566.916 1.804.024 1.804.024 1.804.024 1.804.024

Total Flow kg/hr 3.382.506 3.321.441 2.822.842 3.250.000 3.250.000 3.250.000 3.250.000

Total Flow cum/hr 0.4153845 0.3616315 1.470.023 0.3552321 4.026.669 9.781.966 1.477.861

Temperature C 7.555.421 9.952.220 2.910.346 1.031.990 1.519.166 1.519.166 2.194.948

Pressure bar 1.000.000 1.000.000 5.000.000 5.000.000 5.000.000 5.000.000 5.000.000

Vapor Frac 0 0 1.000.000 0 0.3137519 0.7664397 1.000.000

Liquid Frac 1.000.000 1.000.000 0 1.000.000 0.6862481 0.2335603 0

Solid Frac 0 0 0 0 0 0 0

Enthalpy kcal/mol -6.304.131 -6.687.378 -5.555.851 -6.680.871 -6.295.169 -5.884.585 -5.616.571

Enthalpy kcal/kg -2.699.209 -3.710.180 -3.083.966 -3.708.447 -3.494.350 -3.266.441 -3.117.671

Enthalpy Gcal/hr -0.9130092 -1.232.315 -0.870555 -1.205.245 -1.135.664 -1.061.593 -1.013.243

Entropy cal/mol-K -4.219.837 -3.487.468 -8.529.182 -3.468.275 -2.554.886 -1.588.957 -9.679.756

Entropy cal/gm-K -1.806.787 -1.934.859 -0.4734415 -1.925.185 -1.418.177 -0.8820052 -0.5373081

Density kmol/cum 3.486.580 5.095.649 0.1065912 5.078.439 0.448019 0.1844235 0.12207

Density kg/cum 8.143.074 9.184.603 1.920.271 9.148.949 8.071.188 3.322.440 2.199.125

Average MW 2.335.548 1.802.440 1.801.528 1.801.528 1.801.528 1.801.528 1.801.528

Liq Vol 60F cum/hr 0.3842844 0.3328829 0.2828283 0.3256264 0.3256264 0.3256264 0.3256264

Cost $/hr

72

CW-G-EN CW-G-IN CW-G-OUT CW-GO CW-H2-1 CW-H2-2 CW-H2-3

WW-GAS-1 SPLIT3 WW-H2-2 WW-H2-3 WW-H2-4

TURB1 WW-GAS-5 WW-D2 WW-H2-1 WW-H2-2 WW-H2-3

VAPOR LIQUID VAPOR LIQUID LIQUID MIXED MIXED

Substream: MIXED

Mass Flow kg/hr

CO2 0 0 0 0 0 0 0

O2 0 0 0 0 0 0 0

N2 0 0 0 0 0 0 0

H2O 2.822.842 3.250.000 3.250.000 1.787.500 1.621.380 1.621.380 1.621.380

CO 0 0 0 0 0 0 0

METHANOL 0 0 0 0 0 0 0

H2 0 0 0 0 0 0 0

METHANE 0 0 0 0 0 0 0

DME 0 0 0 0 0 0 0

Liq Vol 60F cum/hr

CO2 0 0 0 0 0 0 0

O2 0 0 0 0 0 0 0

N2 0 0 0 0 0 0 0

H2O 0.2828283 0.3256264 0.3256264 0.1790945 0.1624505 0.1624505 0.1624505

CO 0 0 0 0 0 0 0

METHANOL 0 0 0 0 0 0 0

H2 0 0 0 0 0 0 0

METHANE 0 0 0 0 0 0 0

DME 0 0 0 0 0 0 0


Total Flow kg/hr 2.822.842 3.250.000 3.250.000 1.787.500 1.621.380 1.621.380 1.621.380

Total Flow cum/hr 5.677.799 0.3254012 1.696.983 0.2030094 0.1746689 1.047.517 3.010.116

Temperature C 1.626.700 2.000.000 2.925.393 1.347.313 9.052.701 1.519.166 1.519.166

Pressure bar 1.000.000 5.000.000 5.000.000 5.000.000 5.000.000 5.000.000 5.000.000

Vapor Frac 1.000.000 0 1.000.000 0 0 0.1621855 0.4716145

Liquid Frac 0 1.000.000 0 1.000.000 1.000.000 0.8378145 0.5283855

Solid Frac 0 0 0 0 0 0 0


Enthalpy kcal/kg -3.143.931 -3.794.037 -3.083.249 -3.672.812 -3.722.162 -3.570.657 -3.414.873

Enthalpy Gcal/hr -0.8874822 -1.233.062 -1.002.056 -0.6565152 -0.6035039 -0.5789392 -0.5536806

Entropy cal/mol-K -7.503.847 -3.926.295 -8.506.318 -3.309.252 -3.533.849 -2.878.293 -2.218.045

Entropy cal/gm-K -0.4165268 -2.179.425 -0.4721724 -1.836.914 -1.961.584 -1.597.696 -1.231.202

Density kmol/cum 0.0275972 5.544.000 0.1063077 4.887.524 5.152.620 0.8591769 0.2989927

Density kg/cum 0.4971719 9.987.672 1.915.163 8.805.011 9.282.589 1.547.831 5.386.437

Average MW 1.801.528 1.801.528 1.801.528 1.801.528 1.801.528 1.801.528 1.801.528

Liq Vol 60F cum/hr 0.2828283 0.3256264 0.3256264 0.1790945 0.1624505 0.1624505 0.1624505

Cost $/hr 3.250.000

73

CW-H2-4 CW-H2-IN CW-H2OUT CW-OUT-2 CW-OUT-3 CW-REB-1 CW-REB-2

WW-H2-5 WW-H2-1 SPLIT1 MIX-CW-O MIX-CW-O WW-D1 WW-D2

WW-H2-4 WW-H2-5 SPLIT1 SPLIT3 SPLIT1 SPLIT3

MIXED LIQUID VAPOR VAPOR VAPOR VAPOR VAPOR

Substream: MIXED

Mass Flow kg/hr

CO2 0 0 0 0 0 0 0

O2 0 0 0 0 0 0 0

N2 0 0 0 0 0 0 0

H2O 1.621.380 1.621.380 1.621.380 1.360.342 1.462.500 2.610.377 1.787.500

CO 0 0 0 0 0 0 0

METHANOL 0 0 0 0 0 0 0

H2 0 0 0 0 0 0 0

METHANE 0 0 0 0 0 0 0

DME 0 0 0 0 0 0 0

Liq Vol 60F cum/hr

CO2 0 0 0 0 0 0 0

O2 0 0 0 0 0 0 0

N2 0 0 0 0 0 0 0

H2O 0.1624505 0.1624505 0.1624505 0.1362964 0.1465319 0.026154 0.1790945

CO 0 0 0 0 0 0 0

METHANOL 0 0 0 0 0 0 0

H2 0 0 0 0 0 0 0

METHANE 0 0 0 0 0 0 0

DME 0 0 0 0 0 0 0


Total Flow kg/hr 1.621.380 1.621.380 1.621.380 1.360.342 1.462.500 2.610.377 1.787.500

Total Flow cum/hr 4.229.084 0.1623381 8.419.272 7.063.792 7.636.425 1.355.480 9.333.408

Temperature C 1.519.166 2.000.000 2.894.160 2.894.160 2.925.393 2.894.160 2.925.393

Pressure bar 5.000.000 5.000.000 5.000.000 5.000.000 5.000.000 5.000.000 5.000.000

Vapor Frac 0.6638005 0 1.000.000 1.000.000 1.000.000 1.000.000 1.000.000

Liquid Frac 0.3361995 1.000.000 0 0 0 0 0

Solid Frac 0 0 0 0 0 0 0


Enthalpy kcal/kg -3.318.115 -3.794.037 -3.084.737 -3.084.737 -3.083.249 -3.084.737 -3.083.249

Enthalpy Gcal/hr -0.5379926 -0.6151576 -0.5001531 -0.4196298 -0.4509252 -0.0805232 -0.5511308

Entropy cal/mol-K -1.807.965 -3.926.295 -8.553.831 -8.553.831 -8.506.318 -8.553.831 -8.506.318

Entropy cal/gm-K -1.003.573 -2.179.425 -0.4748098 -0.4748098 -0.4721724 -0.4748098 -0.4721724

Density kmol/cum 0.2128127 5.544.000 0.1068979 0.1068979 0.1063077 0.1068979 0.1063077

Density kg/cum 3.833.880 9.987.672 1.925.796 1.925.796 1.915.163 1.925.796 1.915.163

Average MW 1.801.528 1.801.528 1.801.528 1.801.528 1.801.528 1.801.528 1.801.528

Liq Vol 60F cum/hr 0.1624505 0.1624505 0.1624505 0.1362964 0.1465319 0.026154 0.1790945

Cost $/hr 1.621.380

74

CW-REBOU G-4BI G-4BO G-16BI G-16BO G-64BI G-64BO

WW-GAS-1 GAS-CMP2 WW-GAS-2 GAS-CMP3 WW-GAS-3 GAS-CMP4

WW-D1 GAS-CMP1 WW-GAS-1 GAS-CMP2 WW-GAS-2 GAS-CMP3 WW-GAS-3

LIQUID VAPOR VAPOR VAPOR VAPOR VAPOR VAPOR

Substream: MIXED

Mass Flow kg/hr

CO2 0 2.375.700 2.375.700 2.375.700 2.375.700 2.375.700 2.375.700

O2 0 1.375.200 1.375.200 1.375.200 1.375.200 1.375.200 1.375.200

N2 0 6.982.300 6.982.300 6.982.300 6.982.300 6.982.300 6.982.300

H2O 2.610.377 4.667.000 4.667.000 4.667.000 4.667.000 4.667.000 4.667.000

CO 0 0 0 0 0 0 0

METHANOL 0 0 0 0 0 0 0

H2 0 0 0 0 0 0 0

METHANE 0 0 0 0 0 0 0

DME 0 0 0 0 0 0 0

Liq Vol 60F cum/hr

CO2 0 0.2891112 0.2891112 0.2891112 0.2891112 0.2891112 0.2891112

O2 0 0.2301733 0.2301733 0.2301733 0.2301733 0.2301733 0.2301733

N2 0 1.334.917 1.334.917 1.334.917 1.334.917 1.334.917 1.334.917

H2O 0.026154 0.0467599 0.0467599 0.0467599 0.0467599 0.0467599 0.0467599

CO 0 0 0 0 0 0 0

METHANOL 0 0 0 0 0 0 0

H2 0 0 0 0 0 0 0

METHANE 0 0 0 0 0 0 0

DME 0 0 0 0 0 0 0


Total Flow kg/hr 2.610.377 1.119.990 1.119.990 1.119.990 1.119.990 1.119.990 1.119.990

Total Flow cum/hr 0.0283415 4.027.030 3.272.914 1.298.826 8.375.651 3.318.923 2.093.913

Temperature C 9.739.902 2.474.985 1.500.000 3.985.428 1.600.000 4.134.075 1.600.000

Pressure bar 5.000.000 4.000.000 4.000.000 1.600.000 1.600.000 6.400.000 6.400.000

Vapor Frac 0 1.000.000 1.000.000 1.000.000 1.000.000 1.000.000 1.000.000

Liquid Frac 1.000.000 0 0 0 0 0 0

Solid Frac 0 0 0 0 0 0 0


Enthalpy kcal/kg -3.714.763 -5.308.463 -5.556.829 -4.910.381 -5.531.651 -4.870.302 -5.531.651

Enthalpy Gcal/hr -0.0969693 -0.5945426 -0.6223593 -0.5499577 -0.6195394 -0.545469 -0.6195394

Entropy cal/mol-K -3.498.148 2.779.931 1.190.873 2.045.793 -1.385.086 -0.5295409 -4.138.050

Entropy cal/gm-K -1.941.767 0.0923619 0.0395662 0.0679705 -0.0460188 -0.0175937 -0.1374848

Density kmol/cum 5.112.559 0.0924034 0.1136942 0.2864983 0.4442776 1.121.181 1.777.110

Density kg/cum 9.210.418 2.781.181 3.421.996 8.623.095 1.337.197 3.374.559 5.348.790

Average MW 1.801.528 3.009.824 3.009.824 3.009.824 3.009.824 3.009.824 3.009.824

Liq Vol 60F cum/hr 0.026154 1.900.961 1.900.961 1.900.961 1.900.961 1.900.961 1.900.961

Cost $/hr

75

G-256BI G-256BO G-330BI GAS-BC GAS-IN GAS-S1 GAS-S2

WW-GAS-4 GAS-CMP5 WW-GAS-5 GAS-CMP1 R-STOIC MEMBRSEP

GAS-CMP4 WW-GAS-4 GAS-CMP5 WW-GAS-5 SEP-1 SEP2

VAPOR VAPOR VAPOR VAPOR VAPOR VAPOR VAPOR

Substream: MIXED

Mass Flow kg/hr

CO2 2.375.700 2.375.700 2.375.700 2.375.700 2.375.700 9.540.324 1.588.742

O2 1.375.200 1.375.200 1.375.200 1.375.200 1.375.200 1.356.302 1.887.316

N2 6.982.300 6.982.300 6.982.300 6.982.300 6.982.300 6.905.419 7.680.222

H2O 4.667.000 4.667.000 4.667.000 4.667.000 4.667.000 0.4973201 0.0136719

CO 0 0 0 0 0 1.987.147 0.0238308

METHANOL 0 0 0 0 0 3.252.236 0.4026001

H2 0 0 0 0 0 7.707.370 0.0586936

METHANE 0 0 0 0 0 0.2438557 7.26E+02

DME 0 0 0 0 0 0.0585269 0.0611791

Liq Vol 60F cum/hr

CO2 0.2891112 0.2891112 0.2891112 0.2891112 0.2891112 0.0116101 1.93E+02

O2 0.2301733 0.2301733 0.2301733 0.2301733 0.2301733 0.2270102 3.16E+02

N2 1.334.917 1.334.917 1.334.917 1.334.917 1.334.917 1.320.218 0.0146835

H2O 0.0467599 0.0467599 0.0467599 0.0467599 0.0467599 4.98E+01 1.37E+00

CO 0 0 0 0 0 3.80E+02 4.56E+00

METHANOL 0 0 0 0 0 4.09E+02 5.07E+01

H2 0 0 0 0 0 2.047.690 1.56E+02

METHANE 0 0 0 0 0 8.14E+01 2.43E+00

DME 0 0 0 0 0 8.92E+00 9.33E+00

Total Flow kmol/hr 3.721.114 3.721.114 3.721.114 3.721.114 3.721.114 6.755.552 0.414313

Total Flow kg/hr 1.119.990 1.119.990 1.119.990 1.119.990 1.119.990 9.188.252 1.172.352

Total Flow cum/hr 8.297.307 6.322.466 5.358.135 9.997.799 4.998.454 3.349.291 9.409.294

Temperature C 4.134.075 2.500.000 2.983.646 5.000.000 2.600.000 2.500.000 0

Pressure bar 2.560.000 2.560.000 3.300.000 1.000.000 3.300.000 5.000.000 1.000.000

Vapor Frac 1.000.000 1.000.000 1.000.000 1.000.000 1.000.000 1.000.000 1.000.000

Liquid Frac 0 0 0 0 0 0 0

Solid Frac 0 0 0 0 0 0 0

Enthalpy kcal/mol -1.465.875 -1.595.810 -1.557.962 -1.747.198 -1.588.026 -0.4298614 -1.014.628

Enthalpy kcal/kg -4.870.302 -5.302.005 -5.176.256 -5.804.984 -5.276.142 -3.160.505 -3.585.728

Enthalpy Gcal/hr -0.545469 -0.5938192 -0.5797355 -0.6501524 -0.5909226 -0.0290395 -4.20E+01

Entropy cal/mol-K -3.282.505 -5.441.713 -5.254.065 1.931.310 -5.798.557 -5.978.097 0.5304919

Entropy cal/gm-K -0.1090597 -0.1807984 -0.1745639 0.0641668 -0.1926543 -0.4395324 0.0187477

Density kmol/cum 4.484.725 5.885.542 6.944.794 0.0372193 7.444.530 2.017.009 0.0440323

Density kg/cum 1.349.823 1.771.445 2.090.261 1.120.237 2.240.673 2.743.342 1.245.951

Average MW 3.009.824 3.009.824 3.009.824 3.009.824 3.009.824 1.360.104 2.829.629

Liq Vol 60F cum/hr 1.900.961 1.900.961 1.900.961 1.900.961 1.900.961 3.615.824 0.0220187

Cost $/hr 0

76

H2-4BXI H2-4BXO H2-16BXI H2-16BXO H2-50-XO H2-50BXI H2-200BI

WW-H2-1 H2-CMP-2 WW-H2-2 H2-CMP-3 MIX-H2 WW-H2-3 WW-H2-4

H2-CMP-1 WW-H2-1 H2-CMP-2 WW-H2-2 WW-H2-3 H2-CMP-3 H2-CMP-4

VAPOR VAPOR VAPOR VAPOR VAPOR VAPOR VAPOR

Substream: MIXED

Mass Flow kg/hr

CO2 0 0 0 0 0 0 0

O2 0 0 0 0 0 0 0

N2 0 0 0 0 0 0 0

H2O 0 0 0 0 0 0 0

CO 0 0 0 0 0 0 0

METHANOL 0 0 0 0 0 0 0

H2 3.168.304 3.168.304 3.168.304 3.168.304 3.168.304 3.168.304 1.079.860

METHANE 0 0 0 0 0 0 0

DME 0 0 0 0 0 0 0

Liq Vol 60F cum/hr

CO2 0 0 0 0 0 0 0

O2 0 0 0 0 0 0 0

N2 0 0 0 0 0 0 0

H2O 0 0 0 0 0 0 0

CO 0 0 0 0 0 0 0

METHANOL 0 0 0 0 0 0 0

H2 0.8417534 0.8417534 0.8417534 0.8417534 0.8417534 0.8417534 2.868.967

METHANE 0 0 0 0 0 0 0

DME 0 0 0 0 0 0 0


Total Flow kg/hr 3.168.304 3.168.304 3.168.304 3.168.304 3.168.304 3.168.304 1.079.860

Total Flow cum/hr 1.630.864 1.284.363 5.360.765 3.537.592 1.132.029 1.731.902 1.258.144

Temperature C 2.260.659 1.200.000 3.832.330 1.600.000 1.600.000 3.895.299 2.918.262

Pressure bar 4.000.000 4.000.000 1.600.000 1.600.000 5.000.000 5.000.000 2.000.000

Vapor Frac 1.000.000 1.000.000 1.000.000 1.000.000 1.000.000 1.000.000 1.000.000

Liquid Frac 0 0 0 0 0 0 0

Solid Frac 0 0 0 0 0 0 0

Enthalpy kcal/mol 1.399.823 0.6583398 2.500.568 0.9376027 0.9376027 2.544.715 1.860.251

Enthalpy kcal/kg 6.943.979 3.265.769 1.240.435 4.651.084 4.651.084 1.262.334 9.227.985

Enthalpy Gcal/hr 0.0220006 0.0103469 0.0393007 0.014736 0.014736 0.0399945 0.0996493

Entropy cal/mol-K 0.8589097 -0.8106869 0.0228396 -2.887.200 -5.149.938 -2.172.961

-6.043.344

Entropy cal/gm-K 0.4260718 -0.4021504 0.0113298 -1.432.228 -2.554.685 -1.077.922

-2.997.869

Density kmol/cum 0.0963705 0.1223699 0.2931807 0.4442776 1.388.367 0.907484 4.257.676

Density kg/cum 0.1942715 0.2466829 0.5910172 0.8956103 2.798.782 1.829.379 8.582.963

Average MW 2.015.880 2.015.880 2.015.880 2.015.880 2.015.880 2.015.880 2.015.880

Liq Vol 60F cum/hr 0.8417534 0.8417534 0.8417534 0.8417534 0.8417534 0.8417534 2.868.967

Cost $/hr

77

H2-200BO H2-330BI H2-BC-1 H2-BC-4 H2-IN H20-OUT LIQ-S1

H2CMP-5 WW-H2-5 H2-CMP-1 H2-CMP-4 R-STOIC V2

WW-H2-4 H2CMP-5 MIX-H2 WW-H2-5 SPLIT4 SEP-1

VAPOR VAPOR VAPOR VAPOR VAPOR LIQUID LIQUID

Substream: MIXED

Mass Flow kg/hr

CO2 0 0 0 0 0 3.97E-09 1.625.466

O2 0 0 0 0 0 3.87E-14 1.889.849

N2 0 0 0 0 0 6.20E-14 7.688.098

H2O 0 0 0 0 0 1.366.696 1.391.812

CO 0 0 0 0 0 2.66E-16 0.0238576

METHANOL 0 0 0 0 0 0.1582091 1.586.106

H2 1.079.860 1.079.860 3.168.304 1.079.860 1.079.860 3.91E-22 0.0586969

METHANE 0 0 0 0 0 7.89E-15 7.29E+02

DME 0 0 0 0 0 1.24E-04 0.0658161

Liq Vol 60F cum/hr

CO2 0 0 0 0 0 4.83E-12 1.98E+02

O2 0 0 0 0 0 6.48E-17 3.16E+02

N2 0 0 0 0 0 1.19E-16 0.0146985

H2O 0 0 0 0 0 0.136933 0.1394494

CO 0 0 0 0 0 5.08E-19 4.56E+00

METHANOL 0 0 0 0 0 1.99E+01 0.1996586

H2 2.868.967 2.868.967 0.8417534 2.868.967 2.868.967 1.04E-23 1.56E+02

METHANE 0 0 0 0 0 2.63E-17 2.43E+00

DME 0 0 0 0 0 1.88E-08 1.00E+01


Total Flow kg/hr 1.079.860 1.079.860 3.168.304 1.079.860 1.079.860 1.368.278 3.091.508

Total Flow cum/hr 1.165.001 8.556.861 3.896.047 3.010.897 7.195.575 0.1489752 0.3604416

Temperature C 2.500.000 3.608.634 2.500.000 6.486.496 2.600.000 9.952.220 2.500.000

Pressure bar 2.000.000 3.300.000 1.000.000 5.000.000 3.300.000 1.000.000 5.000.000

Vapor Frac 1.000.000 1.000.000 1.000.000 1.000.000 1.000.000 0 0

Liquid Frac 0 0 0 0 0 1.000.000 1.000.000

Solid Frac 0 0 0 0 0 0 0

Enthalpy kcal/mol 1.567.387 2.343.790 8.90E-12 0.2750922 1.637.403 -6.687.378 -6.230.241

Enthalpy kcal/kg 7.775.199 1.162.663 4.41E-09 1.364.626 8.122.522 -3.710.180 -2.635.593

Enthalpy Gcal/hr 0.0839612 0.1255514 1.40E-13 0.014736 0.0877118 -0.5076558 -0.8147957

Entropy cal/mol-K -6.581.900 -6.230.340 0.0261396 -6.876.615 -7.443.789 -3.487.468 -4.426.373

Entropy cal/gm-K -3.265.026 -3.090.630 0.0129668 -3.411.223 -3.692.575 -1.934.859 -1.872.499

Density kmol/cum 4.598.080 6.260.201 0.0403401 1.779.126 7.444.530 5.095.649 3.628.348

Density kg/cum 9.269.177 1.261.981 0.0813209 3.586.505 1.500.728 9.184.603 8.577.001

Average MW 2.015.880 2.015.880 2.015.880 2.015.880 2.015.880 1.802.440 2.363.886

Liq Vol 60F cum/hr 2.868.967 2.868.967 0.8417534 2.868.967 2.868.967 0.1371322 0.3606776

Cost $/hr 1.584.152

78

LIQ-S2 MEOH-1 MEOH-D1 MEOH-D2 MEOH-S1 MET-OUT1 N2-GAS

MIX-1 V1 DEST-1 DEST-2 SEP-1 MIX-1

SEP2 R-STOIC V2 MIX-1 V1 SPLIT2 MEMBRSEP

LIQUID VAPOR MIXED LIQUID VAPOR LIQUID MIXED

Substream: MIXED

Mass Flow kg/hr

CO2 0.0367239 1.116.579 1.625.466 0.0367239 1.116.579 2.40E-11 9.540.324

O2 2.53E+02 1.375.200 1.889.849 2.53E+02 1.375.200 2.40E-12 1.356.302

N2 7.88E+02 6.982.300 7.688.098 7.88E+02 6.982.300 4.29E-12 6.905.419

H2O 1.360.285 1.396.785 1.391.812 1.391.676 1.396.785 1.378.074 0.4973201

CO 2.67E+00 2.011.005 0.0238576 2.67E+00 2.011.005 1.77E-14 1.987.147

METHANOL 7.528.046 1.618.628 1.586.106 1.582.083 1.618.628 1.506.802 3.252.236

H2 3.28E-01 7.713.240 0.0586969 3.28E-01 7.713.240 2.17E-19 0.770737

METHANE 2.23E+00 0.251142 7.29E+02 2.23E+00 0.251142 2.63E-13 0.2438557

DME 4.64E+02 0.124343 0.0658161 4.64E+02 0.124343 1.10E-04 0.0585269

Liq Vol 60F cum/hr

CO2 4.47E+00 0.0135882 1.98E+02 4.47E+00 0.0135882 2.92E-14 0.0116101

O2 4.24E-01 0.2301733 3.16E+02 4.24E-01 0.2301733 4.02E-15 0.2270102

N2 1.51E+00 1.334.917 0.0146985 1.51E+00 1.334.917 8.20E-15 1.320.218

H2O 1.36E+02 0.1399477 0.1394494 0.1394358 0.1399477 0.1380729 4.98E+01

CO 5.11E-03 3.85E+02 4.56E+00 5.11E-03 3.85E+02 3.38E-17 3.80E+02

METHANOL 9.48E+02 0.2037525 0.1996586 0.1991522 0.2037525 0.1896759 4.09E+02

H2 8.72E-03 2.049.250 1.56E+02 8.72E-03 2.049.250 5.77E-21 0.0204769

METHANE 7.46E-03 8.38E+01 2.43E+00 7.46E-03 8.38E+01 8.77E-16 8.14E+01

DME 7.07E-01 1.90E+01 1.00E+01 7.07E-01 1.90E+01 1.68E-08 8.92E+00

Total Flow kmol/hr 0.3117486 8.063.360 1.307.808 1.266.378 8.063.360 1.235.203 2.970.457

Total Flow kg/hr 8.940.154 1.227.976 3.091.508 2.974.277 1.227.976 2.884.876 8.425.222

Total Flow cum/hr 0.0106077 1.083.125 1.055.777 0.3652334 7.148.629 0.3542735 1.469.869

Temperature C 0 2.600.000 2.378.040 7.377.875 2.600.001 7.555.421 2.500.000

Pressure bar 1.000.000 3.300.000 1.000.000 1.000.000 5.000.000 1.000.000 5.000.000

Vapor Frac 0 1.000.000 0.031641 0 1.000.000 0 0.9979341

Liquid Frac 1.000.000 0 0.968359 1.000.000 0 1.000.000 2.07E+02

Solid Frac 0 0 0 0 0 0 0

Enthalpy kcal/mol -6.031.897 -7.126.076 -6.230.241 -6.297.429 -7.126.076 -6.304.131 -0.9962277

Enthalpy kcal/kg -2.103.359 -4.679.254 -2.635.593 -2.681.299 -4.679.254 -2.699.209 -3.512.372

Enthalpy Gcal/hr -0.0188043 -0.5746012 -0.8147957 -0.7974926 -0.5746012 -0.7786882 -0.0295925

Entropy cal/mol-K -5.378.023 -7.519.193 -4.411.750 -4.246.055 -3.771.767 -4.219.837 -6.744.953

Entropy cal/gm-K -1.875.349 -0.493739 -1.866.313 -1.807.872 -0.2476686 -1.806.787 -0.2378049

Density kmol/cum 2.938.872 7.444.530 1.238.715 3.467.311 1.127.959 3.486.580 2.020.900

Density kg/cum 8.427.934 1.133.734 2.928.182 8.143.497 1.717.778 8.143.074 5.731.956

Average MW 2.867.745 1.522.909 2.363.886 2.348.649 1.522.909 2.335.548 2.836.338

Liq Vol 60F cum/hr 0.0109104 3.976.501 0.3606776 0.3386593 3.976.501 0.3277488 1.588.610

Cost $/hr

79

REC-BHX1 REC-BHX2 REC-IND1 REC-IND2 RECYCLH2 RECYH2IN TOP-D1 TOP-D2

WW-D1 WW-D2 DEST-1 DEST-2 MIX-H2 SEP2

SPLIT2 SPLIT4 WW-D1 WW-D2 MEMBRSEP DEST-1 DEST-2

LIQUID LIQUID MIXED VAPOR VAPOR VAPOR VAPOR LIQUID

Substream: MIXED

Mass Flow kg/hr

CO2 5.10E-10 5.67E-09 5.10E-10 5.67E-09 0 0 1.625.466 0.0367239

O2 3.54E-10 5.52E-14 3.54E-10 5.52E-14 0 0 1.889.849 2.53E+02

N2 7.36E-10 8.85E-14 7.36E-10 8.85E-14 0 0 7.688.098 7.88E+02

H2O 2.377.138 1.950.905 2.377.138 1.950.905 0 0 1.373.957 2.486.287

CO 2.88E-12 3.80E-16 2.88E-12 3.80E-16 0 0 0.0238576 2.67E+00

METHANOL 2.599.210 0.2258373 2.599.210 0.2258373 0 0 7.930.646 1.580.500

H2 3.80E-16 5.58E-22 3.80E-16 5.58E-22 7.630.296 7.630.296 0.0586969 3.28E-01

METHANE 2.00E-11 1.13E-14 2.00E-11 1.13E-14 0 0 7.29E+02 2.23E+00

DME 1.91E-06 1.76E-04 1.91E-06 1.76E-04 0 0 0.0658161 4.64E+02

Liq Vol 60F cum/hr

CO2 6.20E-13 6.90E-12 6.20E-13 6.90E-12 0 0 1.98E+02 4.47E+00

O2 5.92E-13 9.24E-17 5.92E-13 9.24E-17 0 0 3.16E+02 4.24E-01

N2 1.41E-12 1.69E-16 1.41E-12 1.69E-16 0 0 0.0146985 1.51E+00

H2O 0.0238171 0.1954665 0.0238171 0.1954665 0 0 1.38E+02 2.49E+02

CO 5.50E-15 7.26E-19 5.50E-15 7.26E-19 0 0 4.56E+00 5.11E-03

METHANOL 0.0327187 2.84E+01 0.0327187 2.84E+01 0 0 9.98E+02 0.198953

H2 1.01E-17 1.48E-23 1.01E-17 1.48E-23 2.027.213 2.027.213 1.56E+02 8.72E-03

METHANE 6.69E-14 3.76E-17 6.69E-14 3.76E-17 0 0 2.43E+00 7.46E-03

DME 2.91E-09 2.69E-08 2.91E-09 2.69E-08 0 0 1.00E+01 7.07E-01

Total Flow kmol/hr 2.130.696 1.083.622 2.130.696 1.083.622 3.785.094 3.785.094 0.7260616 5.071.874

Total Flow kg/hr 4.976.347 1.953.163 4.976.347 1.953.163 7.630.296 7.630.296 2.066.367 1.605.882

Total Flow cum/hr 0.0611114 0.2126563 5.165.372 3.358.623 1.876.587 1.876.587 1.982.027 0.2141122

Temperature C 7.555.421 9.952.220 8.667.682 9.963.319 2.500.000 2.500.000 5.517.874 6.151.714

Pressure bar 1.000.000 1.000.000 1.000.000 1.000.000 5.000.000 5.000.000 1.000.000 1.000.000

Vapor Frac 0 0 0.8101845 1.000.000 1.000.000 1.000.000 1.000.000 0

Liquid Frac 1.000.000 1.000.000 0.1898155 0 0 0 0 1.000.000

Solid Frac 0 0 0 0 0 0 0 0

Enthalpy kcal/mol -6.304.128

-6.687.378

-5.532.265

-5.714.857

8.90E-12 8.90E-12 -2.708.075

-5.639.500

Enthalpy kcal/kg -2.699.205

-3.710.180

-2.368.721

-3.170.622

4.41E-09 4.41E-09 -9.515.390

-1.781.130

Enthalpy Gcal/hr -0.1343218

-0.7246587

-0.1178758

-0.6192742

3.37E-13 3.37E-13 -0.0196622

-0.2860283

Entropy cal/mol-K -4.219.841

-3.487.468

-2.040.103

-8.785.126

-7.742.526 -7.742.526

-7.939.923

-5.406.790

Entropy cal/gm-K -1.806.787

-1.934.859

-0.8735002

-0.4874018

-3.840.767 -3.840.767

-0.2789859

-1.707.633

Density kmol/cum 3.486.577 5.095.649 0.0412496 0.0322638 2.017.009 2.017.009 0.0366322 2.368.793

Density kg/cum 8.143.074 9.184.603 0.9634054 0.581537 4.066.049 4.066.049 1.042.552 7.500.188

Average MW 2.335.550 1.802.440 2.335.550 1.802.440 2.015.880 2.015.880 2.845.994 3.166.249

Liq Vol 60F cum/hr 0.0565359 0.1957508 0.0565359 0.1957508 2.027.213 2.027.213 0.0329292 0.2015154

Cost $/hr

80

B1 - substitution of reboiler in heat exchanger.

Figure 49: Flowsheet for the substitution of the roboiler in destilation tower 1 in to a heat exchanger.

Figure 50: Design specification and sensivity test on the reboiler substitution.

Due the design Specifications and sensitivity test the exact amount of super-heated steam for the

reboiler was calculated.

This was done in the same way for the reboiler of the second destilation tower.

81

C1 - B.E.U. TEMA sheet

Figure 51: B.E.U. TEMA sheet.

82

C2 - D.E.U. TEMA sheet

Figure 52: D.E.U. TEMA sheet

Improvements on the design of carbon dioxide conversion to ... · Improvements on the design of carbon dioxide conversion to methanol process using Aspen ... the design of carbon

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